Catalyst for producing dimethyl ether, method for producing catalyst and method for producing dimethyl ether

ABSTRACT

A catalyst for producing dimethyl ether which comprises alumina particles having an average size of 200 μm or less and a methanol synthesis catalyst layer formed around the alumina particles. The methanol synthesis catalyst is in an amount of 0.05 to 5 parts by weight to 1 part by weight of the alumina particles. Dimethyl ether is produced by the method of forming a slurry by introducing the catalyst into a solvent and introducing a mixed gas comprising carbon monoxide and hydrogen into the slurry.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a Divisional application of application Ser. No.10/341,478 filed Jan. 13, 2003, now U.S. Pat. No. 6,800,665 which is aContinuation application of application Ser. No. 09/587,153 filed Jun.2, 2000 (U.S. Pat. No. 6,562,306), which is a Divisional application ofapplication Ser. No. 08/847,347 filed Apr. 24, 1997 (U.S. Pat. No.6,147,125).

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a catalyst for dimethyl ether, a methodfor producing the catalyst, and a method for producing dimethyl ether byusing the catalyst

2. Description of the Related Arts

There are several known methods for manufacturing dimethyl etherstarting from a mixed gas of carbon monoxide, carbon dioxide, andhydrogen under the presence of a catalyst suspended in a solvent

For example, JP-A-2-9833 (the term “JP-A-” referred to herein signifies“Unexamined Japanese patent publication”), JP-A-3-181435, JP-A-3-52835,JP-A-4-264046, WO 93/10069 disclose methods for manufacturing dimethylether or a mixture of dimethyl ether and methanol through the contact ofa synthesis gas with a mixture of a methanol synthesis catalyst and amethanol dehydration catalyst suspended in an inert liquid.

The method disclosed in JP-A-2-9833 is a method of direct synthesis ofdimethyl ether from a synthesis gas, which method comprises the step ofcontacting a synthesis gas consisting of hydrogen, carbon monoxide andcarbon dioxide with a solid catalyst, or reacting the synthesis gasreact under the presence of the solid catalyst to conduct catalyticsynthesis of dimethyl ether from the synthesis gas, wherein thesynthesis gas undergoes catalytic action under the presence of the solidcatalyst system, and wherein the solid catalyst is a single catalyst ora mixture of plurality of catalysts which are suspended in a liquidmedium in a three-phase (slurry phase) reactor system, and wherein thethree-phase reactor system comprises at least a single three-phasereactor.

The method disclosed in JP-A-3-181435 is a method for manufacturingdimethyl ether from a mixed gas of carbon monoxide and hydrogen, or amixed gas of carbon monoxide and hydrogen and further containing carbondioxide and/or water vapor, wherein a catalyst is used in a slurry formby suspension thereof in a solvent.

The method disclosed in JP-A-3-52835 is a method of dimethyl ethersynthesis characterized in that a synthesis gas is reacted under thepresence of a solid methanol synthesis catalyst to produce methanol, andthat the produced methanol is reacted under the presence of a soliddehydration catalyst to produce dimethyl ether. According to the method,dimethyl ether is synthesized from a synthesis gas consisting ofhydrogen, carbon monoxide, and carbon dioxide. That is, the synthesisgas is contacted with a solid catalyst system comprising amethanol-synthesizing ingredient and a dehydrating (ether-forming)ingredient, wherein the solid catalyst system is a single catalyst or amixture of plurality of catalysts in a three-phase (liquid phase)reactor system, and wherein the reactor system is controlled to keep theminimum effective methanol rate to at least a level of 1.0 g-mole ofmethanol per 1 kg of catalyst per hour.

The method disclosed in WO 93/10069 is a method for manufacturingdimethyl ether from a mixed gas containing carbon monoxide and either orboth of water and water vapor, or from a mixed gas containing carbonmonoxide and either or both of water and water vapor and furthercontaining carbon dioxide, wherein a catalyst is used in a form ofsolvent slurry, which catalyst is prepared by pulverizing a mixedcatalyst containing at least zinc oxide and, copper oxide or chromiumoxide, and aluminum oxide, by adhering these ingredients together underpressure, and by pulverizing them again to suspend in the solvent.

On the other hand, dimethyl ether is synthesized generally in a fixedbed system. There is a known catalyst for a fixed bed system, whichcatalyst is prepared by depositing a methanol synthesis catalyst onto asupport of metallic oxide such as alumina, then by calcining themtogether. (JP-A-2-280386)

The methods for manufacturing dimethyl ether disclosed in JP-A-2-9833,JP-A-3-52835, JP-A-4-264046, and JP-A-3-181435, however, raise problemssuch that the two kinds of or three kinds of catalysts suspended in asolvent separate from each other in the reactor owing to the differencein specific gravity among the methanol synthesis catalyst, the methanoldehydration catalyst, and the water gas shift catalyst, which induces adistribution in catalyst concentration or deposition of one of thesecatalysts, thus significantly degrading the use efficiency of thecatalysts.

The catalyst disclosed in WO 93/10069 is prepared by integrating theabove-described three kinds of catalysts by means of a mechanicalmethod. These types of catalysts also raise a problem, that during aperiod of use in a slurry state, the catalyst particles separate fromeach other to induce a distribution in catalyst concentration andcatalyst deposition.

SUMMARY OF THE INVENTION

It is an object of the present invention to provide a catalyst suitablefor producing dimethyl ether at a high yield and a method for producingthe catalyst, and to provide a method for producing dimethyl ether at ahigh space time yield.

To attain the object, the present invention provides a first catalystsuitable for producing dimethyl ether, the catalyst comprising:

alumina particles having an average size of 200 μm or less;

a layer comprising a methanol synthesis catalyst, the layer being formedaround each of the alumina particles; and

the methanol synthesis catalyst having a weight ratio of 0.05 to 5 to aweight of the alumina particles.

In the first catalyst, the average size of the alumina particles ispreferably 1 to 100 μm. The average size of 1 to 50 μm is morepreferable.

The methanol synthesis catalyst may comprise copper oxide, zinc oxideand alumina. It is desirable that a weight ratio of the copper oxide:the zinc oxide: the alumina is 1:0.05 to 20:0 to 2. The methanolsynthesis catalyst may comprise zinc oxide, chromium oxide and alumina.It is desirable that a weight ratio of the zinc oxide: the chromiumoxide: the alumina being 1:0.1 to 10:0 to 2.

The first catalyst suitable for producing dimethyl ether is produced bythe following method comprising the steps of:

forming a layer comprising a methanol synthesis catalyst around each ofalumina particles; and

washing the alumina particles, around which the layer was formed, withan acid aqueous solution.

The forming of the layer may comprise:

forming a slurry by introducing the alumina particles into an aqueoussolution containing a metallic salt of active element of the methanolsynthesis catalyst;

heating the slurry; and

neutralizing the heated slurry with a base solution, thereby the activeelement of the methanol synthesis catalyst being deposited around eachof the alumina particles.

The deposition of the active element of the methanol synthesis catalystis preferably carried out at a temperature of 50 to 90° C.

Dimethyl ether is produced by using the first catalyst. A method forproducing dimethyl ether comprising the steps of:

providing the first catalyst suitable for producing dimethyl ether, thecatalyst;

forming a slurry by introducing the catalyst into a solvent; and

introducing a mixed gas comprising carbon monoxide and hydrogen into theslurry.

The present invention provides a second catalyst suitable for producingdimethyl ether, the catalyst comprising:

alumina particles having pores;

deposits which exist inside the pores; and

the deposits comprising copper oxide, zinc oxide, and alumina

It is preferable that the alumina particles have an average size of 200μm or less. A weight ratio of the copper oxide: the zinc oxide: thealumina is preferably 1:0.05 to 20:0 to 2.

The second catalyst suitable for producing dimethyl ether is produced bythe method comprising the steps of:

(a) introducing an alumina having pores into an aqueous solutioncontaining a copper salt, a zinc salt and an aluminum salt forimpregnating the pores with the aqueous solution;

(b) vaporizing the aqueous solution on a surface of the alumina;

(c) contacting the alumina subjected to the step (b) with a solutioncontaining a deposition agent for hydrolyzing the copper salt, the zincsalt and the aluminum salt to copper hydroxide, zinc hydroxide andaluminum hydroxide and for depositing the copper hydroxide, the zinchydroxide and the aluminum hydroxide within the pores of the alumina;

(d) washing the alumina subjected to the step (c); and

(e) calcining the washed alumina.

Dimethyl ether is produced by using the second catalyst A method forproducing dimethyl ether comprising the steps of:

providing the second catalyst suitable for producing dimethyl ether,

forming a slurry by introducing the catalyst into a solvent; and

introducing a mixed gas comprising carbon monoxide and hydrogen into theslurry.

The present invention provides a third catalyst suitable for producingdimethyl ether, the catalyst comprising:

a methanol synthesis catalyst;

a methanol dehydration catalyst; and

a binder for integrating the methanol synthesis catalyst and themethanol dehydration catalyst.

The methanol synthesis catalyst may comprise copper oxide, zinc oxideand alumina A weight ratio of the copper oxide:the zinc oxide:thealumina is preferably 1:0.05 to 20:0 to 2. Also, the methanol synthesiscatalyst may comprise zinc oxide, chromium oxide and alumina. A weightratio of the zinc oxide: the chromium oxide the alumina is preferably1:0.1 to 10:0 to 2. The methanol dehydration catalyst may be at leastone selected from the group consisting of γ-alumina, silica-alumina andzeolite. The third catalyst may further comprise a water gas shiftcatalyst. The water gas shift catalyst may comprise iron oxide andchromium oxide. The binder may be alumina sol or clay.

Dimethyl ether is produced by using the third catalyst. A method forproducing dimethyl ether comprising the steps of:

providing the third catalyst suitable for producing dimethyl ether;

forming a slurry by introducing the catalyst into a solvent; and

introducing a mixed gas comprising carbon monoxide and hydrogen into theslurry.

Further, the present invention provides a method for producing acatalyst suitable for producing dimethyl ether, the method comprisingthe steps of:

(a) preparing a methanol synthesis catalyst, a methanol dehydrationcatalyst, a water gas shift catalyst and a solvent;

b) calculating an A value regarding to the methanol synthesis catalyst,the methanol dehydration catalyst and the water gas shift catalyst usingan average particle size of the catalyst, a particle density of thecatalyst and a density of the solvent, the A value being defied by thefollowing equation:A=D ²(P−S),

-   -   where D denotes the average particle size of the catalyst, (cm),        -   P denotes the particle density of the catalyst, (g/cm³), and        -   S denotes the density of the solvent, (g/cm³),

(c) controlling at least one of the group consisting of the averageparticle size of the catalyst, the particle density of the catalyst andthe density of the solvent to maintain differences in the A valueswithin ±1×10⁻⁶ g/cm among the methanol synthesis catalyst, the methanoldehydration catalyst, and the water gas shift catalyst;

(d) after the step (c), suspending the methanol synthesis catalyst, themethanol dehydration catalyst, and the water gas shift catalyst in thesolvent.

Furthermore, dimethyl ether can be produced by the method comprising thesteps of:

providing a mixed gas containing carbon monoxide and at least oneselected from the group of hydrogen and water vapor;

contacting the mixed gas with a first catalyst consisting essentially ofa methanol synthesis catalyst, a dehydration catalyst and a water gasshift catalyst; and

contacting the mixed gas, which contacted with the first catalyst, witha second catalyst consisting essentially of at least one selected fromthe group of a dehydration catalyst and a water gas shift catalyst.

Moreover, dimethyl ether can be produced by the method comprising thesteps of:

(a) reacting a raw material gas containing carbon monoxide and hydrogenin the presence of a catalyst to produce a reaction gas includingdimethyl ether, carbon dioxide, carbon monoxide and hydrogen;

(b) separating the reaction gas into the carbon monoxide and thehydrogen, and the dimethyl ether and the carbon dioxide;

(c) recycling the carbon monoxide and the hydrogen which were separatedfrom the reaction gas in the step (b);

(d) removing the carbon dioxide from the dimethyl ether and the carbondioxide of the step (b) to gain the dimethyl ether; and

(e) recycling the dimethyl ether which was gained in the step (d) to thestep (b).

Further, dimethyl ether can be produced by the method comprising thesteps of:

preparing a slurry which is produced by dispersing a dimethyl ethersynthesis catalyst into a medium oil;

contacting a raw material gas containing carbon monoxide and hydrogenwith the slurry to produce a product gas containing a vaporized mediumoil;

cooling the product gas to condense the vaporized medium oil;

obtaining dimethyl ether from the product gas from which the vaporizedmedium oil was condensed;

removing a catalyst-deactivation ingredient from the condensed mediumoil; and

recycling the medium oil, from which the catalyst-deactivationingredient was-removed, to the step of preparing the slurry.

Further, the present invention provides an apparatus for producingdimethyl ether comprising:

a slurry-bed reactor filled with a dimethyl ether synthesis catalyst anda medium oil therefor;

a condenser for condensing a vaporized medium oil discharged from thereactor;

an adsorber for removing a catalyst-deactivation ingredient from themedium oil condensed in the condenser, and

recycle means for recycling the medium oil to the slurry-bed reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic representation showing the apparatus for producinga dimethyl ether used in the embodiment 5.

FIG. 2 is a graph showing reaction equilibria of H₂, CO, methanol, CO₂and H₂O.

FIG. 3 is a schematic representation showing the apparatus for producingdimethyl ether used in the embodiment 6.

FIG. 4 is an enlarged view of the non-reacted gas separator in FIG. 3indicating the connection of pipelines.

FIG. 5 is a schematic representation showing the apparatus for producingdimethyl ether used in the embodiment 7.

FIG. 6 is another schematic representation showing the apparatus forproducing dimethyl ether used in the embodiment 7.

FIG. 7 is a graph showing the change of dissolved sulfur concentrationwith time in the case that the apparatus of FIG. 6 was used.

FIG. 8 is a graph showing the change in conversion of carbon monoxidewith time in the case that the apparatus of FIG. 6 was used.

FIG. 9 is a schematic side view of a reactor of the embodiment 8.

FIG. 10 is a part-enlarged view of the reactor of FIG. 9.

FIG. 11 is a schematic representation showing the apparatus forproducing dimethyl ether into which the reactor of the embodiment 8 isinstalled.

FIG. 12 is a schematic side view of a reactor of the embodiment 9.

FIG. 13 is a schematic side view of another reactor of the embodiment 9.

FIG. 14 is a schematic representation showing the apparatus forproducing dimethyl ether into which the reactor of the embodiment 9 isinstalled.

FIG. 15 is a graph showing reaction equilibria of H₂, CO, methanol, CO₂and H₂O.

FIG. 16 is a schematic representation showing the apparatus forproducing dimethyl ether into which the reactor of the embodiment 10 isinstalled.

DESCRIPTION OF THE EMBODIMENT Embodiment 1

The catalyst-according to the present invention comprises aluminaparticles and a methanol synthesis catalyst layer formed around each ofthe alumina particles. Alumina functions as a methanol-dehydrationcatalyst, and alumina in use as an ordinary catalyst may be appliedwithout further processing. A preferable size of alumina particles isfine, preferably 200 μm or less of average particle size, morepreferably in an approximate range of from 1 to 100 μm, and mostpreferably in an approximate range of from 1 to 50 μm. To prepare thepreferred average particle size, alumina may be pulverized as needed.

Methanol synthesis catalyst may be copper oxide-zinc oxide-aluminasystem and zinc oxide-chromium oxide-alumina system. A preferable mixingratio of individual ingredients of copper oxide, zinc oxide, and aluminais: in an approximate range of from 0.05 to 20 wt.parts of zinc oxide to1 wt.parts of copper oxide, more preferably in an approximate range offrom 0.1 to 5 wt.parts; in an approximate range of from 0 to 2 wt.partsof alumina, more preferably in an approximate range of from 0 to 1wt.parts. A preferable mixing ratio of individual ingredients of zincoxide, chromium oxide, and alumina is: in an approximate range of from0.1 to 10 of chromium oxide to 1 wt.parts of zinc oxide, more preferablyin an approximate range of from 0.5 to 5; and in an approximate range offrom 0 to 2 wt.parts of alumina, more preferably in an approximate rangeof from 0 to 1 wt.parts.

A preferable ratio of the methanol synthesis catalyst layer formedaround alumina particle is in a range of from 0.05 to 5 by wt.parts to 1wt.parts of alumina, more preferably in a range of from 0.1 to 3wt.parts, and most preferably in a range of from 0.5 to 2 wt.parts.

The catalyst achieves 30% or higher CO conversion, or normally in anapproximate range of from 35 to 60%, and particularly in an approximaterange of from 45 to 55%, and achieves 20% or higher dimethyl etheryield, or normally in an approximate range of from 25 to 45%, andparticularly in an approximate range of from 35 to 45%. A preferredparticle size of the catalyst is small as far as possible within a rangethat no agglomeration problem occurs. A preferable average particle sizeof the catalyst is 200 μm or less, more preferably in an approximaterange of from 1 to 100 μm, and most preferably in an approximate rangeof from 1 to 50 μm.

The method for manufacturing the catalyst according to the presentinvention is characterized in the steps of: forming a methanol synthesiscatalyst around each of alumina particles; and washing the catalystlayer with an acid aqueous solution. For manufacturing the catalyst,powdered alumina is charged into an aqueous solution containing ametallic salt of active ingredient of methanol synthesis catalyst, forexample an aqueous solution of copper salt, zinc salt, and aluminumsalt, to prepare a slurry. Copper salt, zinc salt, and aluminum salt maybe either inorganic salt or organic salt if only the salt is a watersoluble salt. Nevertheless, a salt that likely generates hydroxide inwater is not suitable. As for the sat of copper and of zinc (andchromium), nitrate, carbonate, organic acid salt are applicable. For thesalt of halide and of aluminum, nitrate, carbonate, and organic salt areapplicable. A preferable concentration of each ingredient is in anapproximate range of from 0.1 to 3 mole/liter.

Thus prepared alumina slurry is then heated, and a base solution isadded dropwise to the heated slurry to neutralize to deposit the activeingredient of the methanol synthesis catalyst around each of aluminaparticles. The alumina particle may be coated by copper salt, zinc salt,and aluminum salt separately at need. A preferable temperature of theslurry during the period of deposition is in a range of from 50 to 90°C., more preferably in a range of from 60 to 85° C. Any kind of base isapplicable if only it can neutralize the acid in the slurry. Theneutralization is to deposit copper, zinc, and aluminum, and apreferable pH value is in an approximate range of from 6 to 12, morepreferably in an approximate range of from 7 to 10. Afterneutralization, the slurry is allowed to stand for an appropriate timeor is subjected to mild agitation for aging to sufficiently develop thedeposits.

The alumina particles on each of which the deposit is formed areseparated from liquid. Thus separated solid alumina particles are washedwith warm water. Normally, succeeding drying and calcining treatmentprovides a catalyst configured by alumina coated by a methanol synthesiscatalyst. With the catalyst according to the present invention, however,methanol which is generated on the methanol synthesis catalyst migratesonto the alumina which is a methanol-dehydration catalyst, where themethanol undergoes dehydration and condensation by the action of acidactive centers on the alumina to yield dimethyl ether, which mechanismis described later. In the deposit-forming operation described above,alumina contacts with base solution, and the acid active centers onalumina vanish. To recover the acid active centers on alumina, themethod for manufacturing catalyst according to the present inventioninvolves washing the alumina particles with an acid aqueous solutionafter forming the above-described deposit. For washing, the aluminaparticles may be suspended in an acid aqueous solution. Applicable acidfor washing is either inorganic acid or organic acid. A preferable acidincludes nitric acid, hydrochloric acid, and acetic acid, and morepreferably nitric acid and hydrochloric acid. The concentration of acidfor acid washing is in an approximate range of from 0.1 to 5 mole/liter,more preferably in an approximate range of from 0.5 to 2 mole/liter. Thetemperature for washing may be room temperature, and a warm temperaturemay also be applicable. The period for washing may be in an approximaterange of from 10 to 30 min.

The next step is washing the alumina particles and deposits thoroughlywith ion-exchanged water or the like to remove acid and base ions,followed by drying and calcining. The calcining may be carried out inair. The calcining temperature may be a temperature level that the metalhydroxide in the catalyst ingredients for methanol-synthesis isconverted to metal oxide. For instance, a preferable calcining conditionis preferably carried out at a temperature range of from 250 to 400° C.for a period of from 1 to 10 hours.

The above-described catalyst is used in a state of slurry of a solvent.The amount of catalyst in the solvent depends on the kind of thesolvent, the reaction conditions, and other variables. Normally, theamount of catalyst is in a range of from 1 to 50 wt. % to the amount ofthe solvent.

The kind of solvent used for synthesizing dimethyl ether according tothe present invention is arbitrarily selected if only the solvent is ina liquid phase under the reaction conditions. Examples of the solventare hydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture.

Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation are also applicable as the solvent.

By passing a mixed gas of carbon monoxide and hydrogen through thusprepared slurry of catalyst with solvent, dimethyl ether is obtained ata high yield. An applicable range of the molar mixing ratio of hydrogento carbon monoxide (H₂/CO) is wide, for instance, in a range of from 20to 0.1 as (H₂/CO), and more preferably from 10 to 0.2.

According to the reaction system, the mixed gas does not directlycontact with catalyst only after being dissolved into a solvent.Consequently, catalytic reaction system, but the carbon monoxide andhydrogen contact with catalyst only after being dissolved into asolvent. Consequently, selection of an adequate kind of solvent takinginto account of the solubility of carbon monoxide and hydrogen to thesolvent establishes a constant composition of carbon monoxide andhydrogen in the solvent independent of the gas composition, and sustainsthe supply of the mixed gas at an established composition to thecatalyst surface.

In the case of a mixed gas with a significantly low ratio of (H₂/CO),for example, 0.1 or less, or in the case of solely carbon monoxidewithout containing hydrogen, it is necessary to separately supply steamto convert a part of the carbon monoxide into hydrogen and carbondioxide within the reactor.

Since solvent exists between the raw material gases and the catalyst,the gas composition does not necessarily agree with the composition onthe catalyst surface. Therefore, it is acceptable that the mixed gas ofcarbon monoxide and hydrogen, or solely carbon monoxide gas includes arelatively high concentration of carbon dioxide (in a range of from 20to 50%)

The manufacturing method according to the present inventionsignificantly reduces the effect of ingredients that may act as acatalyst-poison on the catalyst compared with the gas-solid contactcatalyst system. Examples of the catalyst-poisoning ingredients whichmay exist in the raw material gas are sulfur compounds such as hydrogensulfide, cyan compounds such as hydrogen cyanide, and chlorine compoundssuch as hydrogen chloride. Even when the catalyst activity decreases asa result of poisoning, the productivity of the total reactor system ismaintained at a constant level by withdrawing the slurry from thereactor and by charging fresh slurry containing catalyst of highactivity to the reactor.

The reaction heat is recovered in a form of medium pressure steam usinga cooling coil installed inside of the reactor while passing hot waterthrough the cooling coil. The cooling system controls the reactiontemperature at an arbitrary level.

A preferable reaction temperature is in a range of from 150 to 400° C.,and particularly preferable in a range of from 200 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide.

A preferable reaction pressure is in a range of from 10 to 300 kg/cm²,and particularly preferable in a range of from 15 to 150 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires a special design of thereactor and is uneconomical because of the need of a large amount ofenergy for pressurizing the system.

A preferable space velocity (charge rate of mixed gas per 1 g ofcatalyst under standard condition) is in a range of from 100 to 50000ml/g.h, and particularly preferable from 500 to 30000 ml/g·h. The spacevelocity above 50000 ml/g·h degrades the conversion of carbon monoxide,and that below 100 ml/g·h is uneconomical because of the need of anexcessively large reactor.

The catalyst for manufacturing dimethyl ether according to the presentinvention comprises alumina particles with a methanol synthesis catalystlayer formed around each of the alumina particles. The method formanufacturing the catalyst is characterized in the steps of: forming acatalyst active ingredients for synthesizing methanol around each ofalumina particles; and washing the catalyst active ingredients with anacid aqueous solution. Since every catalyst ingredient has a size of amolecular level and since each of the ingredients chemically adsorbs byeach other, they do not separate during the reaction period. Inaddition, a very small distance of adjacent active ingredients allowsthe reaction cycle described below to proceed promptly, thus improvingthe yield of dimethyl ether. That is, the sequent order of the reactionbegins with the yielding of methanol from carbon monoxide and hydrogenon the methanol synthesis catalyst, then the produced methanol migratesonto the alumina inside of the catalyst, where the methanol undergoesdehydration and condensation on the acid active centers on alumina, thusyielding dimethyl ether and water. Furthermore, the water migrates ontothe methanol synthesis catalyst, where the water reacts with carbonmonoxide to yield carbon dioxide and hydrogen. The reaction follows thereaction formulae given below.CO+2H₂→CH₃OH2CH₃OH→CH₃OCH₃+H₂OCO+H₂O→CO₃+H₂

The method for manufacturing dimethyl ether according to the presentinvention significantly increases the yield of dimethyl ether by usingthe catalyst which comprises alumina particles and methanol synthesiscatalyst layer formed around each of the alumina particles in a state ofslurry with a solvent. The method is free from problems such as pluggingof catalyst and mechanical strength of catalyst, is designed to readilyabsorb the reaction heat through a cooling pipe or the like, and isdesigned to conduct withdrawing and filling of catalyst easily.

EXAMPLE

I. Preparation of Catalyst

Examples 1, 5 through 8

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution,100 g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The preparedslurry was heated to around 80° C. and held at the temperature.Separately, about 1.4 kg of sodium carbonate (Na₂CO₃) was dissolved intoabout 1 liter of ion-exchanged water, which solution was then heated toabout 80° C. The solution was added dropwise to the slurry until theslurry reached to pH 8.0. After completing the dropwise addition, theslurry was allowed to stand for about 1 hour for aging. Then, the slurrywas filtered, and the cake was rinsed by ion-exchanged water at about80° C. until sodium ion and nitric acid ion are not detected anymore.After the rinse, the resultant alumina particles were suspended in about1 liter of aqueous solution of 1 mole/liter nitric acid, and were washedat room temperature for 20 min. The suspension was filtered to recoverthe alumina particles. The alumina particles were further rinsed withion-exchanged water until no acid was detected anymore. The aluminaparticles were then dried at 120° C. for 24 hours followed by calciningthereof in air at 350° C. for 5 hours. The calcined alumina particleswere classified to recover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

Example 2

Each of 37.1 g of copper nitrate (Cu(NO₃)₂.3H₂O), 23.4 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 10.3 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution,100 g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The preparedslurry was heated to around 80° C. and held at the temperature.Separately, about 0.3 kg of sodium carbonate (Na₂CO₃) was dissolved intoabout 1 liter of ion-exchanged water, which solution was then heated toabout 80° C. The solution was added dropwise to the slurry until theslurry reached to pH 8.0. After completing the dropwise addition, theslurry was allowed to stand for about 1 hour for aging. Then, the slurrywas filtered, and the cake was rinsed by ion-exchanged water at about80° C. until sodium ion and nitric acid ion were not detected anymore.After the rinse, the resultant alumina particles were suspended in about1 liter of aqueous solution of 1 mole/liter nitric acid, and were washedfollowing the procedure applied in Example 1. The alumina particles werethen dried at 120° C. for 24 hours, followed by calcining thereof in airat 350° C. for 5 hours. The calcined alumina particles were classifiedto recover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=20:11:69 (by weight).

Example 3

Each of 92.6 g of copper nitrate (Cu(NO₃)₂.3H₂O), 58.5 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 25.5 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution,100 g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The preparedslurry was heated to around 80° C. and held at the temperature.Separately, about 0.7 kg of sodium carbonate (Na₂CO₃) was dissolved intoabout 1 liter of ion-exchanged water, which solution was then heated toabout 80° C. The solution was added dropwise to the slurry until theslurry reached to pH 8.0. After completing the dropwise addition, theslurry was allowed to stand for about 1 hour for aging. Then, the slurrywas filtered, and the cake was rinsed by ion-exchanged water at about80° C. until sodium ion and nitric acid ion were not detected anymore.After the rinse, the resultant alumina particles were suspended in about1 liter of aqueous solution of 1 mole/liter nitric acid, and were washedfollowing the procedure applied in Example 1. The alumina particles werethen dried at 120° C. for 24 hours, followed by calcining thereof in airat 350° C. for 5 hours. The calcined alumina particles were classifiedto recover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

Example 4

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃. 9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution, 50g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The preparedslurry was heated to around 80° C. and held at the temperature.Separately, about 1.4 kg of sodium carbonate (Na₂CO₃) was dissolved intoabout 1 liter of ion-exchanged water, which solution was then heated toabout 80° C. The solution was added dropwise to the slurry until theslurry reached to pH 8.0. After completing the dropwise addition, theslurry was allowed to stand for about 1 hr for aging. Then, the slurrywas filtered, and the cake was rinsed by ion-exchanged water at about80° C. until sodium ion and nitric acid ion were not detected anymore.After the rinse, the resultant alumina particles were washed by about. 1liter of aqueous solution of 1 mole/liter nitric acid following theprocedure applied in Example 1. The alumina particles were then dried at120° C. for 24 hours followed by calcining thereof in air at 350° C. for5 hours. The calcined alumina particles were classified to recover 120μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=41:21:38 (by weight).

Example 9

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 40° C.

Example 10

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 50° C.

Example 11

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 60° C.

Example 12

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 70° C.

Example 13

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 90° C.

Example 14

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 95° C.

Comparative Example 1

A catalyst was prepared following the procedure applied in Example 1except that the temperature of slurry was set to 20° C.

Comparative Example 2

A catalyst was prepared following the procedure applied in Example 1except that the cake was rinsed only with ion-exchanged water and notrinsed with aqueous solution of nitric acid.

Comparative Example 3

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃. 9H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about1.4 kg of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of these solutions were added dropwise toabout 3 liters of ion-exchanged water in a stainless steel vesselcontrolled at about 80° C., for a period of about 2 hours whilemaintaining the pH value of the mixture to 8.0±0.5. After completing thedropwise addition, the mixture was allowed to stand for about 1 hour foraging. When, during the processing period, pH value came outside of arange of 8.0±0.5, an aqueous solution of about 1 mole/liter nitric acidor of about 1 mole/liter sodium carbonate was added dropwise to themixture to sustain the pH value in a range of 8.0 ±0.5. Then, thegenerated precipitate was filtered, and the cake was rinsed byion-exchanged water until nitric acid ion is not detected in thefiltrate anymore. The cake was then dried at 120° C. for 24 hoursfollowed by calcining thereof in air at 350° C. for 5 hours. A 50 gportion of the calcined cake was powdered in a ball mill along with 50 gof γ-alumina (N612, Nikki Kagaku Co.) for about 3 hours. The powdermixture was calcined in air at 450° C. for 3 hours. The calcined mixturewas further powdered to about 120 μm or finer size as the targetcatalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

II. Method for Activating Catalyst and Reaction Method

A 24 g of n-hexadecane (31.1 ml) was charged to a bubble-tower reactorhaving 2 cm of inside diameter and 2 m of height, and 3.6 g of each ofthe above-described catalyst powders was added to make the contents ofthe reactor in a suspended state. In Examples 5 through 8, the catalystof Example 1 was used. A mixed gas of hydrogen, carbon monoxide, andnitrogen (at a molar ratio H₂:CO:N₂ of 1:1:9) was introduced to passthrough the bubble-tower at a flow rate of about 300 ml/min. Whileflowing the mixed gas through the bubble-tower, the temperature in thebubble-tower was gradually raised from room temperature to 220° C.within a period of several hours. At the same time, the concentration ofnitrogen in the mixed gas was gradually reduced to a final level ofzero. Then, the reaction system was held at 220° C. for about 3 hours toactivate the catalyst.

The reaction was conducted at a specified temperature and pressure whileintroducing the mixed gas of hydrogen, carbon monoxide, and carbondioxide at a molar ratio of H₂/CO/CO₂=47.5/47.5/5.0 and at a flow rateof 336 ml/min. (converted at a condition of normal temperature andpressure).

The obtained reaction products and non-reacted substances were analyzedby gas chromatography.

III. Reaction Conditions and Experimental Results

The reaction conditions and experimental results are shown in Tables 1through 5.

TABLE 1 Example 1 Example 2 Example 3 Example 4 Condition Temperature (°C.) 280 280 280 280 Pressure (kg/cm²-G) 30 30 30 30 Reaction COconversion (%) 47.7 37.8 40.5 48.3 Result Yield Dimethyl ether 36.5 25.328.4 35.7 (C-mol %) Methanol 2.0 0.7 1.1 2.4 Hydrocarbons 0.3 0.2 0.30.1 CO₂ 8.9 11.6 10.7 10.1 Dimethyl ether space time yield 998 692 777976 (g/kg-cat · h)

TABLE 2 Example 5 Example 6 Example 7 Example 8 Condition Temperature (°C.) 250 300 280 280 Pressure (kg/cm²-G) 30 30 20 50 Reaction COconversion (%) 40.3 45.1 38.4 55.9 Result Yield Dimethyl ether 24.2 33.826.5 43.6 (C-mol %) Methanol 3.8 1.7 1.5 2.2 Hydrocarbons 0.1 2.8 0.40.8 CO₂ 12.2 6.8 10.0 9.3 Dimethyl ether space time yield 662 924 7251192 (g/kg-cat · h)

TABLE 3 Example 9 Example 10 Example 11 Example 12 Condition Temperature(° C.) 280 280 280 280 Pressure (kg/cm²-G) 30 30 30 30 Reaction COconversion (%) 38.8 40.2 46.9 47.8 Result Yield Dimethyl ether 28.5 31.136.1 35.9 (C-mol %) Methanol 3.1 2.4 1.8 2.2 Hydrocarbons 0.1 0.1 0.20.2 CO₂ 7.1 6.6 8.8 9.5 Dimethyl ether space time yield 779 850 987 982(g/kg-cat · h)

TABLE 4 Comparative Comparative Example 13 Example 14 Example 1 Example2 Condition Temperature (° C.) 280 280 280 280 Pressure (kg/cm²-G) 30 3030 30 Reaction CO conversion (%) 41.9 40.0 30.1 18.4 Result YieldDimethyl ether 32.1 29.3 17.5 6.5 (C-mol %) Methanol 1.7 2.0 5.7 6.5Hydrocarbons 0.1 0.2 0.1 0.1 CO₂ 8.0 8.5 6.8 1.3 Dimethyl ether spacetime yield 878 801 478 178 (g/kg-cat · h)

TABLE 5 Reaction result in Reaction result in Example 1 after 100Example 3 after 100 hours of continuous hours of continuous reactionreaction Condition Temperature (° C.) 280 280 Pressure (kg/cm²-G) 30 30Reaction CO conversion (%) 41.9 27.3 Result Yield Dimethyl ether 36.313.8 (C-mol %) Methanol 1.8 9.8 Hydrocarbons 0.2 0.5 CO₂ 8.8 3.2Dimethyl ether 992 377 space time yield (g/kg-cat · h)

The catalyst for manufacturing dimethyl ether according to the presentinvention provides effects of preventing separation of individualcatalyst ingredients from each other during reaction, of assuring smoothprogress of reaction cycle, and of achieving high dimethyl ether yieldowing to the configuration thereof comprising alumina particles andmethanol synthesis catalyst layer formed around each of the aluminaparticles.

The method for manufacturing dimethyl ether according to the presentinvention uses a slurry of solvent with a catalyst comprising aluminaparticles and methanol synthesis catalyst layer formed around each ofthe alumina particles, so the method provides effects of achieving highspace time yield of dimethyl ether, of being free from problems ofplugging of catalyst and of mechanical strength of catalyst, of easinessfor removing reaction heat and for controlling reaction heat, ofassuring wide application range of the ratio of carbon monoxide tohydrogen, of progress of reaction under the presence of highconcentration of carbon dioxide, and of less influence of impurities andcatalyst poisons.

Embodiment 2

The catalyst according to the present invention comprises an aluminahaving micropores with deposits of copper oxide, zinc oxide, and aluminatherein. Alumina functions as a methanol-dehydration catalyst, andalumina in use as an ordinary catalyst may be applied without furtherprocessing. A preferable size of alumina particles is a fine size,preferably 200 μm or less of average particle size, more preferably inan approximate range of from 1 to 100 μm, and most preferably in anapproximate range of from 1 to 50 μm. To prepare the preferred averageparticle size, alumina may be pulverized as needed.

The ratio of copper oxide, zinc oxide, and alumina which are depositedon the surface of micropores in alumina is in an approximate range offrom 0.05 to 20 wt.parts of zinc oxide to 1 wt.parts of copper oxide,more preferably in an approximate range of from 0.1 to 5 wt.parts; in anapproximate range of from 0 to 2 wt.parts of alumina to 1 wt.parts ofcopper oxide, more preferably in an approximate range of from 0 to 0.5wt.parts. A preferable deposition amount is in an approximate range offrom 0.05 to 5 wt.parts of the sum of copper oxide, zinc oxide, andalumina to 1 wt.parts of the alumina particles being deposited thereon,and more preferably in an approximate range of from 0.1 to 3 wt.parts,and most preferably in an approximate range of from 0.5 to 2 wt.parts.

The catalyst achieves 25% or higher CO conversion, or normally in anapproximate range of from 30 to 50%, and particularly in an approximaterange of from 40 to 50%, and achieves 20% or higher dimethyl etheryield, or normally in an approximate range of from 25 to 35%, andparticularly in an approximate range of from 30 to 35%. A preferredparticle size of the catalyst is small as far as possible within a rangethat no agglomeration problem occurs. A preferable average particle sizeof the catalyst is 200 μm or less, more preferably in an approximaterange of from 1 to 100 μm, and most preferably in an approximate rangeof from 1 to 50 μm.

The method for manufacturing the catalyst according to the presentinvention is characterized in the steps of: depositing copper oxide,zinc oxide, and alumina onto the surface of micropores in alumina usinga base solution; then calcining the alumina and the deposits. Formanufacturing the catalyst, granulated alumina particles or aluminapowder is impregnated with an aqueous solution containing adequate kindsof copper salt, zinc salt, and aluminum salt. The kind of copper salt,zinc salt, and aluminum salt may be either an inorganic salt or anorganic salt if only the salt is soluble in water. Nevertheless, a saltthat likely generates hydroxide in water by hydrolysis is not suitable.As for the salt of copper and of zinc, nitrate, carbonate, organic acidsalt are applicable. For the salt of halide and of aluminum, nitrate,carbonate, and organic salt are applicable. A preferable concentrationof each ingredient is in an approximate range of from 0.1 to 3mole/liter.

After the aqueous solution of the salts fully penetrated into themicropores in alumina, excess amount of aqueous solution is removed asneeded, then the remaining aqueous solution is vaporized to dry. At thepoint that the surface water on the alumina is vaporized off, or at thepoint that the total amount of applied aqueous solution of the saltsfilled the micropores in alumina, the alumina is brought into contactwith a solution of the deposition agent. The deposition agent is a base,which reacts with an inorganic portion that structures the copper salt,the zinc salt, and the aluminum salt to form a water-soluble salt, whilethe deposition agent itself is able to be emitted by thermaldecomposition during the succeeding calcining step. A preferabledeposition agent includes ammonia, urea, and an organic base, andparticularly ammonia is preferred. A base which cannot be thermallydecomposed, such as sodium hydroxide, potassium hydroxide, and sodiumcarbonate, is not favorable because that type of base remains inmicropores in alumina even after sufficient washing with water tointerfere with the catalyst activity. A preferable concentration of thesolution of deposition agent is in an approximate range of from 0.5 to10 mole/liter. With the addition of solution of deposition agent, coppersalt, zinc salt, and aluminum salt are hydrolyzed within the microporesin alumina by the base as a deposition agent, thus depositing thosesalts onto the surface of the micropores. The deposition of copper,zinc, and aluminum may be conducted separately, at need. Applicabledeposition temperature is around 80° C. at the maximum, and roomtemperature is preferred. As for the pH value, the pH in the microporesis important, though the pH within the micropores cannot be measured.The pH value in the solution of deposition agent is always at alkaliside (pH>12). The deposits are then fully washed with ion-exchangedwater or the like to remove base ions and inorganic ions, and thealumina with deposits is dried and calcined. The calcining may beconducted in air. The temperature of calcining is preferably at atemperature that copper hydroxide, zinc hydroxide, and aluminumhydroxide are converted into copper oxide, zinc oxide, and alumina,respectively, and that base is thermally decomposed to emit, or, forexample, in a range of from 250 to 400° C. for a period of 1 to 10hours.

Thus prepared catalysts are used in a state of slurry with a solventafter classifying to remove a portion of the catalysts which excessivelyenlarged their size caused by deposition. The amount of catalyst in thesolvent depends on the kind of the solvent, the reaction conditions, andother variables. Normally, the amount of catalyst is in a range of from1 to 50 wt. % to the amount of the solvent.

The kind of solvent used for synthesizing dimethyl ether according tothe present invention is arbitrarily selected if only the solvent is ina liquid phase under the reaction condition. Examples of the solvent arehydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture.

Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation are also applicable as the solvent.

By passing a mixed gas of carbon monoxide and hydrogen through thusprepared slurry of catalyst with solvent, dimethyl ether is obtained ata high yield. An applicable range of molar mixing ratio of hydrogen tocarbon monoxide (H₂/CO) is wide, for instance, in a range of from 20 to0.1 as (H₂/CO), and more preferably from 10 to 0.2. According to thereaction system, the mixed gas does not directly contact with thecatalyst, which direct contact occurs in a gas-solid catalytic reactionsystem, but the carbon monoxide and hydrogen contact with catalyst onlyafter dissolved into a solvent. Consequently, selection of an adequatekind of solvent taking into account of the solubility of carbon monoxideand hydrogen to the solvent establishes a constant composition of carbonmonoxide and hydrogen in the solvent independent of the gas composition,and sustains the supply of the mixed gas at an established compositionto the catalyst surface.

In the case of a mixed gas with a significantly low ratio of (H₂/CO),for example, 0.1 or less, or in the case of solely carbon monoxidewithout containing hydrogen, it is necessary to separately supply steamto convert a part of the carbon monoxide into hydrogen and carbondioxide within the reactor.

Since solvent exists between the raw material gases and the catalyst,the gas composition does not necessarily agree with the composition onthe catalyst surface. Therefore, it is acceptable that the mixed gas ofcarbon monoxide and hydrogen, or solely carbon monoxide gas includes arelatively high concentration of carbon dioxide (in a range of from 20to 50%)

The manufacturing method according to the present inventionsignificantly reduces the effect of ingredients that may act as acatalyst-poison on the catalyst compared with the gas-solid contactcatalyst system. Examples of the catalyst-poisoning ingredients whichmay exist in the raw material gas are sulfur compounds such as hydrogensulfide, cyan compounds such as hydrogen cyanide, and chlorine compoundssuch as hydrogen chloride. Even when the catalyst activity is decreasedas a result of poisoning, the productivity of the total reactor systemis maintained at a constant level by withdrawing the slurry from thereactor and by charging fresh slurry containing catalyst of highactivity to the reactor.

The reaction heat is recovered in a form of medium pressure steam usinga cooling coil installed inside of the reactor while passing hot waterthrough the cooling coil. The cooling system controls the reactiontemperature at an arbitrary level.

A preferable reaction temperature is in a range of from 150 to 400° C.,and particularly preferable in a range of from 200 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide.

A preferable reaction pressure is in a range of from 10 to 300 kg/cm²,and particularly preferable in a range of from 15 to 150 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires a special design of thereactor and is uneconomical because of the need of a large amount ofenergy for pressurizing the system.

A preferable space velocity (charge rate of mixed gas per 1 g ofcatalyst under standard condition) is in a range of from 100 to 50000ml/g.h, and particularly preferable from 500 to 30000 ml/g·h. The spacevelocity above 50000 ml/g·h degrades the conversion of carbon monoxide,and that below 100 μl/g·h is uneconomical because of the need of anexcessively large reactor.

The catalyst for manufacturing dimethyl ether according to the presentinvention is prepared by depositing copper oxide, zinc oxide and aluminaonto the surface of micropores in alumina Individual catalystingredients do not separate from each other during the reaction process.Therefore, the reaction cycle proceeds smoothly, and a high yield ofdimethyl ether is attained.

The method for manufacturing dimethyl ether according to the presentinvention significantly increases the yield of dimethyl ether by usingthe catalyst in a state of a slurry with a solvent, which catalyst isprepared by depositing copper oxide, zinc oxide, and alumina on thesurface of micropores in alumina, The method is free from problems suchas plugging of catalyst and mechanical strength of catalyst, is designedto readily absorb the reaction heat through a cooling pipe or the like,and is designed to conduct withdrawing and filing of catalyst easily.

EXAMPLE

I. Preparation of Catalyst

Examples 1, 5 through 8

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution,100 g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The water of themixture was removed by vaporizing in a water-bath. The resulted materialwas put into about 1 liter of about 5 mole/liter aqueous ammonia. Themixture was held for about 1 hour. Then the mixture was further washeduntil no ammonium ion nor nitric acid ion was detected anymore. Themixture was dried at 120° C. for 24 hours. followed by calcined in airat 350° C. for 5 hours. The calcined particles were classified torecover 120 μm or finer ones as the target catalyst

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

Example 2

Each of 37.1 g of copper nitrate (Cu(NO₃)₂.3H₂O), 23.4 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 10.3 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 200 ml of ion-exchanged water. To the solution, 100g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The water of themixture was removed by vaporizing in a water-bath. The resultantmaterial was put into about 1 liter of about 1 mole/liter aqueousammonia. The mixture was held for about 1 hour. Then the mixture wasfurther washed until no ammonium ion nor nitric acid ion was detectedanymore. The mixture was dried at 120° C. for 24 hours. followed bycalcined in air at 350° C. for 5 hours. The calcined particles wereclassified to recover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=20:11:69 (by weight).

Example 3

Each of 92.6 g of copper nitrate (Cu O₃)₂.3H₂O), 58.5 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 25.5 g of aluminum nitrate (Al(NO₃)₃. 9H₂O) weredissolved into about 500 ml of ion-exchanged water. To the solution, 100g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The water of themixture was removed by vaporizing in a water-bath. The resultantmaterial was put into about 1 liter of about 2.5 mole/liter aqueousammonia. The mixture was held for about 1 hour. Then the mixture wasfurther washed until no ammonium ion nor nitric acid ion was detectedanymore. The mixture was dried at 120° C. for 24 hours. followed bycalcined in air at 350° C. for 5 hours. The calcined particles wereclassified to recover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

Example 4

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃. 9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution, 50g of fine powder of γ-alumina (N612, Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The water of themixture was removed by vaporizing in a water-bath. The resulted materialwas put into about 1 liter of about 5 mole/liter aqueous ammonia. Themixture was held for about 1 hour. Then the mixture was further washeduntil no ammonium ion nor nitric acid ion was detected anymore. Themixture was dried at 120° C. for 24 hours. followed by calcined in airat 350° C. for 5 hours. The calcined particles were classified torecover 120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=41:21:38 (by weight).

Example 9

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃.9H₂O) weredissolved into about 1 liter of ion-exchanged water. To the solution,100 g of fine powder of γ-alumina (N612 Nikki Kagaku Co.) having anapproximate particle size of 20 μm or less was added. The water of themixture was removed by vaporizing in a water-bath. The resulted materialwas put into about 1 liter of about 10 mole/liter aqueous solution ofurea. The mixture was heated to about 90° C. under agitation. At thepoint that pH value reached to about 8, the mixture was allowed to standfor cooling. Then the mixture was filtered, and the cake was rinsed. Thecake was dried at 120° C. for 24 hours. followed by calcined in air at350° C. for 5 hours. The calcined particles were classified to recover120 μm or finer ones as the target catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:16:53 (by weight).

II. Method for Activating Catalyst and Reaction Method

A 24 g of n-hexadecane (31.1 ml) was charged to a bubble-tower reactorhaving 2 cm of inside diameter and 2 m of height, and 3.6 g of each ofthe above-described catalyst powders was added to make the contents ofthe reactor in a suspended state. In Examples 5 through 8, the catalystof Example 1 was used. A mixed gas of hydrogen, carbon monoxide, andnitrogen (at a molar ratio. H₂:CO:N₂ of 1:1:9) was introduced to passthrough the bubble-tower at a flow rate of about 300 ml/min. Whileflowing the mixed gas through the bubble-tower, the temperature in thebubble-tower was gradually raised from room temperature to 220° C.within a period of several hours. At the same time, the concentration ofnitrogen in the mixed gas was gradually reduced to a final level ofzero. Then, the reaction system was held at 220° C. for about 3 hours.to activate the catalyst.

The reaction was conducted at a specified temperature and pressure whileintroducing the mixed gas of hydrogen, carbon monoxide, and carbondioxide at a molar ratio of H₂/CO/CO₂=47.5/47.5/5.0 and at a flow rateof 336 ml/min. (converted at a condition of normal temperature andpressure).

The obtained reaction products and non-reacted substances were analyzedby gas chromatography.

III. Reaction Conditions and Experimental Results

The reaction conditions and experimental results are shown in Tables 6through 8.

TABLE 6 Example 1 Example 2 Example 3 Example 4 Condition Temperature (°C.) 280 280 280 280 Pressure (kg/cm²-G) 30 30 30 30 Reaction COconversion (%) 41.3 27.1 30.3 38.6 Result Yield Dimethyl ether 31.5 20.924.1 27.3 (C-mol %) Methanol 1.8 0.7 1.4 2.0 Hydrocarbons 0.1 0.2 0.20.1 CO₂ 7.9 5.3 4.6 9.2 Dimethyl ether space time yield 861 572 659 747(g/kg-cat · h)

TABLE 7 Example 5 Example 6 Example 7 Example 8 Condition Temperature (°C.) 250 300 280 280 Pressure (kg/cm²-G) 30 30 20 50 Reaction COconversion (%) 29.5 38.4 32.2 45.6 Result Yield Dimethyl ether 23.8 32.424.6 33.2 (C-mol %) Methanol 2.1 1.1 0.8 1.7 Hydrocarbons 0.1 0.6 0.10.5 CO₂ 3.5 4.3 6.7 10.2 Dimethyl ether space time yield 651 886 673 908(g/kg-cat · h)

TABLE 8 Example 9 Condition Temperature (° C.) 280 Pressure (kg/cm²-G)30 Reaction CO conversion (%) 40.8 Result Yield Dimethyl ether 30.7(C-mol %) Methanol 1.7 Hydrocarbons 0.2 CO₂ 8.2 Dimethyl ether 840 spacetime yield (g/kg-cat · h)

The catalyst for manufacturing dimethyl ether according to the presentinvention provides effects of preventing separation of individualcatalyst ingredients from each other during reaction, of assuring smoothprogress of reaction cycle, and of achieving high dimethyl ether yieldowing to the configuration thereof comprising an alumina havingmicropores with deposits of copper oxide, zinc oxide, and aluminatherein.

The method for manufacturing dimethyl ether according to the presentinvention uses a slurry of solvent with a catalyst comprising an aluminahaving micropores with deposits of copper oxide, zinc oxide, and aluminatherein, so the method provides effects of achieving high space timeyield of dimethyl ether, of being free from problems of plugging ofcatalyst and of mechanical strength of catalyst, of easiness forremoving reaction heat and for controlling reaction heat, of assuringwide application range of the ratio of carbon monoxide to hydrogen, ofprogress of reaction under the presence of high concentration of carbondioxide, and of less influence of impurities and catalyst poisons.

Embodiment 3

The catalyst for manufacturing dimethyl ether according to the presentinvention is characterized in that at least a methanol synthesiscatalyst and a methanol-dehydration catalyst are integrated togetherusing a binder.

The method for manufacturing dimethyl ether according to the presentinvention comprises the step of charging a mixed gas of carbon monoxideand hydrogen, or a mixed gas of carbon monoxide, hydrogen, and furthercontaining carbon dioxide and/or water vapor to a slurry of the catalystwith a solvent.

The catalyst according to the present invention is basically a catalystin which a methanol synthesis catalyst, a methanol-dehydration catalyst,and a water gas shift catalyst are integrated together. Since, however,the methanol synthesis catalyst is inherently an excellent water gasshift catalyst, the methanol synthesis catalyst may play a role of thewater gas shift catalyst.

The methanol synthesis catalyst may be copper oxide-zinc oxide-aluminasystem and zinc oxide-chromium oxide-alumina system. A preferable mixingratio of individual ingredients of copper oxide, zinc oxide, and aluminais: in an approximate range of from 0.05 to 20 wt.parts of zinc oxide to1 wt.parts of copper oxide, more preferably in an approximate range offrom 0.1 to 5 wt.parts; in an approximate range of from 0 to 2 wt.partsof alumina, more preferably in an approximate range of from 0 to 1wt.parts. A preferable mixing ratio of individual ingredients of zincoxide, chromium oxide, and alumina is: in an approximate range of from0.1 to 10 of chromium oxide to 1 wt.parts of zinc oxide, more preferablyin an approximate range of from 0.5 to 5; and in an approximate range offrom 0 to 2 wt.parts of alumina, more preferably in an approximate rangeof from 0 to 1 wt.parts. Applicable methanol-dehydration catalystsincludes γ-alumina, silica-alumina, and zeolite. Examples of metallicoxide ingredient in zeolite are oxide of alkali metal such as sodium andpotassium, and oxide of alkali earth metal such as calcium andmagnesium. As described above, the methanol synthesis catalyst maysubstitute for the water gas shift catalyst. Other than the methanolsynthesis catalyst, iron oxide-chromium oxide may substitute for thewater gas shift catalyst. A preferable ratio of chromium oxide to ironoxide is in an approximate range of from 0.1 to 20 wt.parts to 1wt.parts of iron oxide, more preferably in an approximate range of from0.5 to 10.

Each of these methanol synthesis catalyst, methanol-dehydrationcatalyst, and water gas shift catalyst may be manufactured by knownmethods. For example, a water soluble salt of each metallic ingredientis used to prepare an aqueous solution containing these salts. The kindof salt is either inorganic salt or organic salt. Nevertheless, a saltthat likely generates hydroxide in water is not suitable. Nitrate,carbonate, organic acid salt, and halide are applicable as the salt. Apreferable concentration of each ingredient is in an approximate rangeof from 0.1 to 3 mole/liter. A base is added to the aqueous solution toneutralize the solution and to precipitate hydroxide. Then the solid isseparated from liquid, which is then rinsed and dried, followed bycalcining to obtain the target catalyst. Alternatively, a commercialcatalyst may be used.

The mixing ratio of above-described methanol synthesis catalyst,methanol-dehydration catalyst, and water gas shift catalyst is notspecifically limited, and is selected at an adequate ratio correspondingto the kind of each ingredient and the reaction condition. A preferablemixing ratio is often in an approximate range of from 0.5 to 10 wt.partsof methanol-dehydration catalyst per 1 wt.parts of methanol synthesiscatalyst, and in an approximate range of from 0 to 5 wt.parts of watergas shift catalyst per 1 wt.parts of methanol synthesis catalyst.

These catalysts are co-pulverized, or pulverized in a mixed state.Co-pulverization is preferably done to an approximate average particlesize of 200 μm or less, more preferably in an approximate range of from1 to 100 μm, and most preferably in an approximate range of from 1 to 50μm.

Thus prepared catalyst mixture particles are blended with a binder. Thebinder needs to have characteristics of durability to calcining(alteration is acceptable), function to bind individual catalysts tointegrate them after calcining, and not degrading the activity of eachcatalyst. Alumina sol and clay are examples of the binder. A preferableratio of binder to catalyst is in a range of from 0.005 to 1 wt.parts to1 wt.parts of catalyst, more preferably in a range of from 0.01 to 1,and most preferably in a range of from 0.05 to 1.

After the co-pulverization, the mixture is processed by drying andcalcining. The calcining may be carried out at a temperature rangingapproximately from 250 to 500° C. for a period of approximately from 2to 10 hours.

After the calcining, the mixture is pulverized for use.

Preferably, the pulverization is conducted to an average particle sizeof 200 μm or less, more preferably to an approximate range of from 1 to100 μm, and most preferably to an approximate range of from 1 to 50 μm.

Thus prepared catalyst achieves 35% or higher CO conversion, normally inan approximate range of from 40 to 60%, and particularly in anapproximate range of from 50 to 60%, and achieves 25% or higher dimethylether yield, or normally in an approximate range of from 30 to 40%.

A primary characteristic of the method according to the presentinvention is to integrate a methanol synthesis catalyst, amethanol-dehydration catalyst, and a water gas shift catalyst using abinder. The integration is conducted for ensuring swift progress of thereaction cycle described below by making the distance among thesecatalysts very short, so that the catalysts are very close to each otherwithout separating them from each other during the reaction period, thusimproving the yield of dimethyl ether. That is, the reaction proceeds inthe sequent order of: yielding methanol on the methanol synthesiscatalyst from carbon monoxide and hydrogen; moving the methanol onto themethanol-dehydration catalyst to yield dimethyl ether and water bydehydration-condensation; moving the water onto the water gas shiftcatalyst and/or the methanol synthesis catalyst to yield hydrogen byreacting with carbon monoxide. The reaction is expressed by the formulaegiven below.CO+2H₂→CH₃OH2CH₃OH→CH₃OCH₃+H₂OCO+H₂O→CO₃+H₂

The above-described catalyst is used in a state of slurry of a solvent.The amount of catalyst in the solvent depends on the kind of thesolvent, the reaction conditions, and other variables. Normally, theamount of catalyst is in a range of from 1 to 50 wt. % to the amount ofthe solvent.

The kind of solvent used for synthesizing dimethyl ether according tothe present invention is arbitrarily selected if only the solvent is ina liquid phase under the reaction condition. Examples of the solvent arehydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture.

Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation are also applicable as the solvent.

By passing a mixed gas of carbon monoxide and hydrogen through thusprepared-slurry of catalyst with solvent, dimethyl ether is obtained ata high yield. Applicable range of molar mixing ratio of hydrogen tocarbon monoxide (H₂/CO) is wide, for instance, in a range of from 20 to0.1 as (H₂/CO), and more preferably from 10 to 0.2.

According to the reaction system, the mixed gas does not directlycontact with the catalyst, which direct contact occurs in a gas-solidcatalytic reaction system, but the carbon monoxide and hydrogen contactwith catalyst only after being dissolved into a solvent. Consequently,selection of an adequate kind of solvent taking into account of thesolubility of carbon monoxide and hydrogen to the solvent establishes aconstant composition of carbon monoxide and hydrogen in the solventindependent of the gas composition, and sustains the supply of the mixedgas at an established composition to the catalyst surface.

In the case of a mixed gas with a significantly low ratio of (H₂/CO),for example, 0.1 or less, or in the case of solely carbon monoxidewithout containing hydrogen, it is necessary to separately supply steamto convert a part of the carbon monoxide into hydrogen and carbondioxide within the reactor.

Since solvent exists between the raw material gases and the catalyst,the gas composition does not necessarily agree with the composition onthe catalyst surface. Therefore, it is acceptable that the mixed gas ofcarbon monoxide and hydrogen, or solely carbon monoxide gas includes arelatively high concentration of carbon dioxide (in a range of from 20to 50%).

The manufacturing method according to the present inventionsignificantly reduces the effect of ingredients that may act as acatalyst-poison on the catalyst compared with the gas-solid contactcatalyst system. Examples of the catalyst-poisoning ingredients whichmay exist in the raw material gas are sulfur compounds such as hydrogensulfide, cyan compounds such as hydrogen cyanide, and chlorine compoundssuch as hydrogen chloride. Even when the catalyst activity is decreasedas a result of poisoning, the productivity of total reactor system ismaintained at a constant level by withdrawing the slurry from thereactor and by charging fresh slurry containing catalyst of highactivity to the reactor.

The reaction heat is recovered in a form of medium pressure steam usinga cooling coil installed inside of the reactor while passing hot waterthrough the cooling coil. The cooling system controls the reactiontemperature at an arbitrary level.

A preferable reaction temperature is in a range of from 150 to 400° C.,and particularly preferable in a range of from 200 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide.

A preferable reaction pressure is in a range of from 10 to 300 kg/cm²,and particularly preferable in a range of from 15 to 150 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires a special design of thereactor and is uneconomical because of the need of a large amount ofenergy for pressurizing the system.

A preferable space velocity (charge rate of mixed gas per 1 g ofcatalyst under standard condition) is in a range of from 100 to 50000ml/g·h, and particularly preferable from 500 to 30000 ml/g·h. The spacevelocity above 50000 ml/g·h degrades the conversion of carbon monoxide,and that below 100 ml/g·h is uneconomical because of the need of anexcessively large reactor.

The catalyst for manufacturing dimethyl ether according to the presentinvention and a method therefor apply integration of a methanolsynthesis catalyst, a methanol-dehydration catalyst, and a water gasshift catalyst with a binder, so the individual catalyst ingredients donot separate from each other during the reaction period. As a result,the reaction cycle proceeds smoothly and a high dimethyl ether yield isachieved.

The method for manufacturing dimethyl ether according to the presentinvention significantly increases the yield of dimethyl ether by usingthe catalyst which comprises an integrated condition of a methanolsynthesis catalyst, a methanol-dehydration catalyst, and a water gasshift catalyst, in a state of a slurry with a solvent. The method isfree from problems such as plugging of catalyst and mechanical strengthof catalyst, is designed to readily absorb the reaction heat through acooling pipe or the like, and is designed to conduct withdrawing andfilling of catalyst easily.

EXAMPLE

I. Preparation of Catalyst

1) Preparation of CuO—ZnO—Al₂O₃ catalyst

Each of 185 g of copper nitrate (Cu(NO₃)₂.3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂.6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃. 9H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about1.4 kg of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of the solutions were added dropwise to about3 liters of ion-exchanged water in a stainless steel vessel which wascontrolled at about 60° C. within about 2 hours, while maintaining thecontents to pH 7.0±0.5. Then, the contents were allowed to stand forabout 1 hour for aging. When, during the treatment, the pH value wentout from a range of pH 7.0±0.5, either of an aqueous solution of about 1mole/liter nitric acid or an aqueous solution of about 1 mole/litersodium carbonate was added dropwise to keep the range of pH 7.0±0.5. Theresultant precipitate was filtered, and the cake was rinsed byion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours. followed bycalcining thereof in air at 350° C. for 5 hours to obtain the targetmethanol synthesis catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=31:32:7 (by weight).

2) Preparation of CuO—Cr₂O₃—ZnO catalyst

Each of 0.30 kg of copper nitrate (Cu(NO₃)₂.3H₂O), 105 g of chromiumnitrate (Cr(NO₃)₂.3H₂O), and 256 g of zinc nitrate (Zn(NO₃)₂.6H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about130 g of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of the solutions were added dropwise to about3 liters of ion-exchanged water in a stainless steel vessel which wascontrolled at about 60° C. within about 2 hours while maintaining thecontents to pH 8.5±0.5. Then, the contents were allowed to stand forabout 1 hr for aging. When, during the treatment, the pH value went outfrom a range of pH 8.5±0.5, either of an aqueous solution of about 1mole/liter nitric acid or an aqueous solution of about 1 mole/litersodium carbonate was added dropwise to keep the range of pH 8.5±0.5. Theresultant precipitate was filtered, and the cake was rinsed byion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours followed bycalcining thereof in air at 350° C. for 5 hours to obtain the targetmethanol synthesis catalyst.

Analysis of thus obtained catalyst gave the composition asCuO:Cr₂O₃:ZnO=1:2:3 (by weight).

3) Preparation of Cu—Al₂O₃ catalyst

A 15.7 g of copper acetate (Cu(CH₃COO)₂·H₂O) was dissolved into about200 ml of ion-exchanged water. A 95 g of γ-alumina (N612, Nikki KagakuCo.) was further added to the mixture. The mixture was then vaporized todry. The dried material was further dried in air at 120° C. for 24hours, followed by calcining in air at 450° C. for 4 hours. The calcinedmaterial was treated in hydrogen gas stream at 400° C. for 3 hours toobtain a catalyst. Analysis of the catalyst gave the composition asCu:Al₂O₃=5:95 (by weight).

Examples 1, 3, and 4

An 100 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment, 50 g of commercially available alumina (N612, Nikki KagakuCo.), and a 50 g aliquot of CuO—Cr₂O₃—ZnO catalyst prepared in the abovetreatment were pulverized together in a ball mill for about 3 hours to afine powder having about 20 μm or less of particle size. A 50 g ofalumina sol (Alumina sol-520, Nissan Chemical Industries, Ltd.) wasadded to the catalyst powder to uniformly mix them together. The mixturewas dried in air at 120° C. for 24 hours followed by calcining in air at450° C. for 3 hours to unify these ingredients. The calcined materialwas pulverized to 120 μm or smaller particle size to obtain a catalyst.

Example 2

A 200 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment, 50 g of commercially available alumina (N612, Nikki KagakuCo.), and a 50 g aliquot of CuO—Cr₂O₃—ZnO catalyst prepared in the abovetreatment were pulverized together in a ball mill for about 3 hours to afine powder having about 20 μm or less of particle size. A 50 g ofalumina sol (Alumina sol-520, Nissan Chemical Industries, Ltd.) wasadded to the catalyst powder to uniformly mix them together. The mixturewas dried in air at 120° C. for 24 hours followed by calcining in air at450° C. for 3 hours to unify these ingredients. The calcined materialwas pulverized to 120 μm or smaller particle size to obtain a catalyst.

Examples 5, 7, 8, 9, and 10

An 100 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment and a 50 g aliquot of Cu—Al₂O₃ catalyst prepared in the abovetreatment were-pulverized together in a ball mill for about 3 hours to afine powder having about 20 μm or less of particle size. A 50 g ofalumina sol (Alumina sol-520, Nissan Chemical Industries, Ltd.) wasadded to the catalyst powder to uniformly mix them together. The mixturewas dried in air at 120° C. for 24 hours followed by calcining in air at450° C. for 3 hours to unify these ingredients. The calcined materialwas pulverized to 120 μm or smaller particle size to obtain anintegrated catalyst.

Example 6

A 200 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment and a 50 g aliquot of Cu—Al₂O₃ catalyst prepared in the abovetreatment were pulverized together in a ball mill for about 3 hours to afine powder having about 20 μm or less of particle size. A 50 g ofalumina sol (Alumina sol-520, Nissan Chemical Industries, Ltd.) wasadded to the catalyst powder to uniformly mix them together. The mixturewas dried in air at 120° C. for 24 hours, followed by calcining in airat 450° C. for 3 hours to unify these ingredients. The calcined materialwas pulverized to 120 μm or smaller particle size to obtain anintegrated catalyst.

Example 11

An 100 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment, 50 g of commercially available alumina (N612, Nikki KagakuCo.), and a 50 g aliquot of CuO—Cr₂O₃—ZnO catalyst prepared in the abovetreatment were pulverized together in a ball mill for about 3 hours to afine powder having about 20 μm or less of particle size. A 50 g of clay(Kunipia, Kunimine Kogyo Co.) was added to the catalyst powder touniformly mix them together. The mixture was dried in air at 120° C. for24 hours followed by calcining in air at 450° C. for 3 hours to unifythese ingredients. The calcined material was pulverized to 120 μm orsmaller particle size to obtain a catalyst.

Comparative Example

An 100 g aliquot of CuO—ZnO—Al₂O₃ catalyst prepared in the abovetreatment, 50 g of commercially available alumina (N612, Nikki KagakuCo.), and a 50 g aliquot of CuO—Cr₂O₃—ZnO catalyst prepared in the abovetreatment were pulverized together in a ball mill for about 0.3 hours toa fine powder having about 20 μm or less of particle size. The mixturewas calcined in air at 450° C. for 3 hours. The calcined material waspulverized to 120 μm or smaller particle size to obtain a catalyst

II. Method for Activating Catalyst and Reaction Method

A 24 g of n-hexadecane (31.1 ml) was charged to a bubble-tower reactorhaving 2 cm of inside diameter and 2 m of height, and 3.6 g of each ofthe above-described catalyst powders was added to make the contents ofthe reactor in a suspended state. A mixed gas of hydrogen, carbonmonoxide, and nitrogen (at a molar ratio H₂:CO:N₂ of 1:1:9) wasintroduced to pass through the bubble-tower at a flow rate of about 300ml/min. While flowing the mixed gas through the bubble-tower, thetemperature in the bubble-tower was gradually raised from roomtemperature to 220° C. within a period of several hours. At the sametime, the concentration of nitrogen in the mixed gas was graduallyreduced to a final level of zero. Then, the reaction system was held at220° C. for about 3 hours to activate the catalyst.

The reaction was conducted at a specified temperature and pressure whileintroducing the mixed gas of hydrogen, carbon monoxide, and carbondioxide at a molar ratio of H₂/CO/CO, 47.5/47.5/5.0 and at a flow rateof 336 ml/min. (converted at a condition of normal temperature andpressure).

The obtained reaction products and non-reacted substances were analyzedby gas chromatography.

III. Reaction Conditions and Experimental Results

The reaction conditions and experimental results are shown in Tables 9and 10.

TABLE 9 Example 1 Example 2 Example 3 Example 4 Example 5 Example 6Condition Temperature (° C.) 280 280 250 300 280 280 Pressure (kg/cm²-G)30 30 30 30 30 30 Reaction CO conversion (%) 51.7 48.3 45.4 47.1 52.951.3 Result Yield Dimethyl 35.2 31.4 27.7 29.6 35.8 33.9 ether (C-mol %)Methanol 2.1 5.3 6.9 1.8 2.4 2.6 Hydrocarbons 0.5 0.6 0.4 3.0 0.8 1.3CO₂ 14.0 11.1 10.4 12.7 13.9 13.5 Dimethyl ether space time 963 859 758809 979 927 yield (-g/kg-cat · h)

TABLE 10 Example Example Comparative Example 7 Example 8 Example 9 10 11Example Condition Temperature(° C.) 260 300 280 280 280 280 Pressure(kg/cm²-G) 30 30 20 50 30 30 Reaction CO conversion (%) 46.5 50.4 38.857.1 44.8 32.4 Result Yield Dimethyl 25.1 33.9 27.2 37.1 29.5 22.9 ether(C-mol %) Methanol 5.9 2.2 3.1 2.9 3.1 1.1 Hydrocarbons 0.3 3.3 0.5 1.70.6 0.2 CO₂ 15.2 11.0 8.0 15.4 11.6 8.2 Dimethyl ether space time 686927 744 1015 807 626 yield (g/kg-cat · h)

The catalyst for manufacturing dimethyl ether according to the presentinvention provides effects of preventing separation of individualcatalyst ingredients from each other during reaction, of assuring smoothprogress of reaction cycle, and of achieving high dimethyl ether yieldowing to the configuration thereof comprising a methanol synthesiscatalyst, a methanol-dehydration catalyst, and a water gasshift-catalyst being integrated together using a binder.

The method for manufacturing dimethyl ether according to the presentinvention uses a slurry of solvent with a catalyst comprising a methanolsynthesis catalyst, a methanol-dehydration catalyst, and a water gasshift catalyst being integrated together using a binder, so the methodprovides effects of achieving high space time yield of dimethyl ether,of being free from problems of plugging of catalyst and of mechanicalstrength of catalyst, of easiness for removing reaction heat and forcontrolling reaction heat, of assuring wide application range of theratio of carbon monoxide to hydrogen, of progress of reaction under thepresence of high concentration of carbon dioxide, and of less influenceof impurities and catalyst poisons.

Embodiment 4

The inventors developed a method for manufacturing a catalyst to producedimethyl ether, which method comprises: suspending a methanol synthesiscatalyst, a methanol-dehydration catalyst, and a water gas shiftcatalyst in a solvent while maintaining the difference in an A valuewithin ±1×10⁻⁶ g/cm among the methanol synthesis catalyst, themethanol-dehydration catalyst, and the water gas shift catalyst, whereinthe A is defined by an equation of A=D²(P−S), where D denotes averageparticle size of catalyst concerned, (cm), P denotes particle density ofcatalyst concerned, (g/cm³), and S denotes solvent density, (g/cm³). Theinventors found that the use of thus prepared catalyst in a slurry stateachieved the production of dimethyl ether at a high yield and a highspace time yield from a mixed gas of carbon monoxide and hydrogen, orfurther containing carbon dioxide and/or water vapor. Thus the inventorscompleted the present invention.

According to the present invention, each of the methanol synthesiscatalyst, the methanol-dehydration catalyst, and the water gas shiftcatalyst is prepared under a control of the particle density and theparticle size, then they are mixed together by a physical means. As aresult, the method according to the present invention ensures promptprogress of the reaction described below to improve the yield ofdimethyl ether by making the distance among these catalysts close toeach other while preventing separation thereof during the reactionperiod. That is, the reaction proceeds in sequent order of: yieldingmethanol on the methanol synthesis catalyst from carbon monoxide andhydrogen; moving the methanol onto the methanol-dehydration catalyst toyield dimethyl ether and water by dehydration-condensation; moving thewater onto the water gas shift catalyst and/or the methanol synthesiscatalyst to yield hydrogen by the reaction of the water with carbonmonoxide. The reaction is expressed by the formulae given below.CO+2H₂→CH₃OH2CH₃OH→CH₃OCH₃+H₂OCO+H₂O→CO₃+H₂

The catalyst according to the present invention basically comprises amethanol synthesis catalyst, a methanol-dehydration catalyst, and awater gas shift catalyst. Since, however, the methanol synthesiscatalyst is inherently an excellent water gas shift catalyst, themethanol synthesis catalyst may function also the water gas shiftcatalyst.

Methanol synthesis catalyst may be copper oxide-zinc oxide-aluminasystem and zinc oxide-chromium oxide-alumina system. A preferable mixingratio of individual ingredients of copper oxide, zinc oxide, and aluminais: in an approximate range of from 0.05 to 20 wt.parts of zinc oxide to1 wt.parts of copper oxide, more preferably in an approximate range offrom 0.1 to 5 wt.parts; in an approximate range of from 0 to 2 wt.partsof alumina, more preferably in an approximate range of from 0 to 1wt.parts. A preferable mixing ratio of individual ingredients of zincoxide, chromium oxide, and alumina is: in an approximate range of from0.1 to 10 of chromium oxide to 1 wt.parts of zinc oxide, more preferablyin an approximate range of from 0.5 to 5; and in an approximate range offrom 0 to 2 wt parts of alumina, more preferably in an approximate rangeof from 0 to 1 wt.parts. Applicable methanol-dehydration catalystsincludes γ-alumina, silica-alumina, and zeolite. Examples of metallicoxide ingredient in zeolite are oxide of alkali metal such as sodium andpotassium, and oxide of alkali earth metal such as calcium andmagnesium. Examples of water gas shift catalyst are a copper oxide-zincoxide system and an iron oxide-chromium oxide system. A preferable ratioof copper oxide to zinc oxide is in an approximate range of from 0.1 to20 wt.parts of copper oxide to 1 wt.parts of iron oxide, more preferablyin an approximate range of from 0.5 to 10. A preferable ratio of ironoxide to chromium oxide is in an approximate range of from 0.1 to 20wt.parts of chromium oxide to 1 wt.parts of iron oxide, more preferablyin an approximate range of from 0.5 to 10%. A copper (including copperoxide)-alumina system is a catalyst which functions both as amethanol-dehydration catalyst and as a water gas shift catalyst.

Each of these methanol synthesis catalyst, methanol-dehydrationcatalyst, and water gas shift catalyst may be manufactured by knownmethods. For example, a water soluble salt of each metallic ingredientis used to prepare an aqueous solution containing these salts. The kindof salt is either an inorganic salt or an organic salt. Nevertheless, asalt that likely generates hydroxide in water is not suitable. Nitrate,carbonate, organic acid salt, and halide are applicable as the salt. Apreferable concentration of each ingredient is in an approximate rangeof from 0.1 to 3 mole/l. A base is added to the aqueous solution toneutralize thereof and to precipitate hydroxide. Then the solid isseparated from liquid, which is then rinsed and dried, followed bycalcining to obtain the target catalyst. Alternatively, a commercialcatalyst may be used.

The mixing ratio of the above-described methanol synthesis catalyst,methanol-dehydration catalyst, and water gas shift catalyst is notspecifically limited, and is selected at an adequate ratio correspondingto the kind of each ingredient and the reaction condition. A preferablemixing ratio is often in an approximate range of from 0.5 to 10 wt.partsof methanol-dehydration catalyst per 1 wt.parts of methanol synthesiscatalyst, and in an approximate range of from 0 to 5 wt.parts of watergas shift catalyst per 1 wt.parts of methanol synthesis catalyst.

A preferable difference in value of A that is derived from theabove-described equation among individual catalyst types is within±1×10⁻⁶ g/cm which was given before, and more preferably within ±5×10⁻⁷g/cm. If the difference in A value exceeds ±1×10⁻⁶, then the conversionof carbon monoxide decreases. The method for controlling the value A isnot specifically limited. Generally, however, the control is carried outbased mainly on the average particle size and the particle density ofcatalyst concerned because the density of solvent shows no significantchange. If only the average particle size is kept constant, the particledensity does not show significant change. Accordingly, the control ofaverage particle size is a simple control method. Pulverization by ballmill is a method for controlling the average particle size. Thedetermination of average particle size is done by sieve method (JISZ8801-1982), sedimentation method, or other methods. The determinationof particle density is carried out by picnometer method (JIS R-5201) andbuoyancy method (JIS R6125).

The catalyst is used as a slurry with a solvent. The amount of catalystin the solvent depends on the kind of the solvent, the reactionconditions, and other variables. Normally, the amount of catalyst is ina range of from 1 to 50 wt. % to the amount of the solvent.

The kind of solvent used for synthesizing dimethyl ether according tothe present invention is arbitrarily selected if only the solvent is ina liquid phase under the reaction condition. Examples of the solvent arehydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture.

Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation are also applicable as the solvent.

By passing a mixed gas of carbon monoxide and hydrogen through thusprepared slurry of catalyst with solvent, dimethyl ether is obtained ata high yield. Applicable range of molar mixing ratio of hydrogen tocarbon monoxide (H₂/CO) is wide, for instance, in a range of from 20 to0.1 as (H₂/CO), and more preferably from 10 to 0.2.

According to the reaction system, the mixed gas does not directlycontact with the catalyst, which direct contact occurs in a gas-solidcatalytic reaction system, but the carbon monoxide and hydrogen contactwith catalyst only after dissolved into a solvent. Consequently,selection of an adequate kind of solvent taking into account of thesolubility of carbon monoxide and hydrogen to the solvent establishes aconstant composition of carbon monoxide and hydrogen in the solventindependent of the gas composition, and sustains the supply of the mixedgas at an established composition to the catalyst surface.

In the case of a mixed gas with a significantly low ratio of (H₂/CO),for example, 0.1 or less, or in the case of solely carbon monoxidewithout containing hydrogen, it is necessary to separately supply steamto convert a part of the carbon monoxide into hydrogen and carbondioxide within the reactor.

Since solvent exists between the raw material gases and the catalyst,the gas composition does not necessarily agree with the composition onthe catalyst surface. Therefore, it is acceptable that the mixed gas ofcarbon monoxide and hydrogen, or solely carbon monoxide gas includes arelatively high concentration of carbon dioxide (in a range of from 20to 50%).

The manufacturing method according to the present inventionsignificantly reduces, the effect of ingredients that may act as acatalyst-poison on the catalyst compared with the gas-solid contactcatalyst system. Examples of the catalyst-poisoning ingredients whichmay exist in the raw material gas are sulfur compounds such as hydrogensulfide, cyan compounds such as hydrogen cyanide, and chlorine compoundssuch as hydrogen chloride. Even when the catalyst activity is decreasedas a result of poisoning, the productivity of the total reactor systemis maintained at a constant level by withdrawing the slurry from thereactor and by charging fresh slurry containing catalyst of highactivity to the reactor.

The reaction heat is recovered in a form of medium pressure steam usinga cooling coil installed inside of the reactor while passing hot waterthrough the cooling coil. The cooling system controls the reactiontemperature at an arbitrary level.

A preferable reaction temperature is in a range of from 150 to 400° C.,and particularly preferable in a range of from 200 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide.

A preferable reaction pressure is in a range of from 10 to 300 kg/cm²,and particularly preferable in a range of from 15 to 150 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires special design of reactorand is uneconomical because of the need of a large amount of energy forpressurizing the system.

A preferable space velocity (charge rate of mixed gas per 1 g ofcatalyst under standard condition) is in a range of from 100 to 50000ml/g·h, and particularly preferable from 500 to 30000 ml/g·h. The spacevelocity above 50000 ml/g·h degrades the conversion of carbon monoxide,and that below 100 ml/g·h is uneconomical because of the need of anexcessively large reactor.

EXAMPLE

I. Preparation of Catalyst

1) Preparation of Catalyst [1]

Each of 185 g of copper nitrate (Cu(NO₃)₂3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃9H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about200 g of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of the solutions were added dropwise to about3 liters of ion-exchanged water in a stainless steel vessel which wascontrolled at about 60° C. within about 2 hours while maintaining thecontents to pH 7.0±0.5. Then, the contents were allowed to stand forabout 1 hour for aging. When, during the treatment, the pH value wentout from a range of pH 7.0±0.5, either of an aqueous solution of about 1mole/liter nitric acid or an aqueous solution of about 1 mole/litersodium carbonate was added dropwise to keep the range of pH 7.0±0.5. Theresulted precipitate was filtered, and the cake was rinsed byion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours, followed bycalcining thereof in air at 350° C. for 3 hours to obtain the catalyst[1].

Analysis of thus obtained catalyst [1] gave the composition asCuO:ZnO:Al₂O₃=61:37:7 (by weight).

2) Preparation of Catalyst [2]

Each of 91 g of copper nitrate (Cu(NO₃)₂3H₂O) and 256 g of zinc nitrate(Zn(NO₃)₂6H₂O) were dissolved into about 1 liter of ion-exchanged water.Separately, about 130 g of sodium carbonate (Na₂CO₃) was dissolved intoabout 1 liter of ion-exchanged water. Both of the solutions were addeddropwise to about 3 liters of ion-exchanged water in a stainless steelvessel which was controlled at about 60° C. within about 2 hours, whilemaintaining the contents to pH 8.5±0.5. Then, the contents were allowedto stand for about 1 hr for aging. When, during the treatment, the pHvalue went out from a range of pH 8.5±0.5, either of an aqueous solutionof about 1 mole/l nitric acid or an aqueous solution of about 1mole/liter sodium carbonate was added dropwise to keep the range of pH8.5±0.5. The resultant precipitate was filtered, and the cake was rinsedby ion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours, followed bycalcining thereof in air at 350° C. for 3 hours to obtain the catalyst[2].

Analysis of thus obtained catalyst [2] gave the composition asCuO:ZnO=3:7 (by weight).

3) Preparation of Catalyst [3]

An 100 g of alumina (N612, Nikki Kagaku Co.) was dried in air at 120° C.for 24 hours, followed by calcining in air at 450° C. for 3 hours toobtain the target alumina catalyst [3].

4) Preparation of catalyst [4]

A 15.7 g of copper acetate (Cu(CH₃COO)₂H₂O) was dissolved into about 200ml of ion-exchanged water. A 95 g aliquot of alumina catalyst preparedin the step 3) was further added to the mixture. The mixture was thenvaporized to dry. The dried material was further dried in air at 120° C.for 24 hours, followed by calcining in air at 450° C. for 3 hours. Thecalcined material was treated in hydrogen gas stream at 400° C. for 3hours to obtain the catalyst [4]. Analysis of the catalyst gave thecomposition as Cu:Al₂O₃=5:95 (by weight).

5) Preparation of Catalyst [5]

A 736 g of aluminum nitrate (Al(NO₃)₃9H₂O) was dissolved into about 2liters of ion-exchanged water. Separately, about 350 g of sodiumcarbonate (Na₂CO₃) was dissolved into about 2 liters of ion-exchangedwater. Both of the solutions were added dropwise to about 3 liters ofion-exchanged water in a stainless steel vessel at room temperaturewithin about 2 hours while maintaining the contents to pH 7.5±0.5. Then,the contents were allowed to stand for about 1 hour for aging. When,during the treatment, the pH value went out from a range of pH 7.5±0.5,either of an aqueous solution of about 1 mole/liter nitric acid or anaqueous solution of about 1 mole/liter sodium carbonate was addeddropwise to keep the range of pH 7.5±0.5. The resulted precipitate wasfiltered, and the cake was rinsed by ion-exchanged water until nitricacid ion was not detected anymore. After the rinse, the cake was driedat 120° C. for 24 hours, followed by calcining thereof in air at 350° C.for 3 hours to obtain alumina.

Then, 15.7 g of copper acetate (Cu(CH₃COO)₂H₂O) was dissolved into about200 ml of ion-exchanged water. A 95 g aliquot of the prepared aluminawas added to the mixture, which mixture was then evaporated to dry. Thedried material was further dried in air at 120° C. for 24 hours followedby calcining in air at 450° C. for 4 hours. The calcined material wastreated in hydrogen gas stream at 400° C. for 3 hours. to obtain thecatalyst [5]. Analysis of thus obtained catalyst [5] gave thecomposition as CuO:Al₂O₃=5:95 (by weight).

Example 1

The catalyst [1] was pulverized in a ball mill to fine powder having anaverage particle size of 16.9 μm. The catalyst [2] was pulverized in aball mill to fine powder having an average particle size of 15.6 μm. Thecatalyst [3] was pulverized in a ball mill to fine powder having anaverage particle size of 15.5 μm. A 2.4 g aliquot of the fine powder ofcatalyst [1], an 1.2 g aliquot of the fine powder of catalyst [2], andan 1.2 g aliquot of the fine powder of catalyst [3] were physicallymixed together.

Example 2

The catalyst [1] was pulverized in a ball mill to fine powder having anaverage particle size of 16.9 μm. The catalyst [4] was pulverized in aball mill to fine powder having an average particle size of 15.2 μm. A2.4 g aliquot of the fine powder of catalyst [1] and an 1.2 g aliquot ofthe fine powder of catalyst [4] were physically mixed together.

Example 3

A 2.4 g aliquot of the fine powder of catalyst [1] having an averageparticle size of 14.4 μm and an 1.2 g aliquot of the fine powder ofcatalyst [4] having an average particle size of 12.9 μm were physicallymixed together.

Example 4

A 2.4 g aliquot of the fine powder of catalyst [1] having an averageparticle size of 16.9 μm and an 1.2 g aliquot of the fine powder ofcatalyst [5] having an average particle size of 18.4 μm were physicallymixed together.

Comparative Example 1

Catalysts were mixed conforming to the procedure of Example 1 exceptthat the average particle size of the catalyst [2] was 20.1 μm and thatthe average particle size of the catalyst [3] was 18.5 μm.

Comparative Example 2

Catalysts were mixed conforming to the procedure of Example 2 exceptthat the average particle size of the catalyst [4] was 12.9 μm.

II. Method for Activating Catalyst and Reaction Method

A 24 g of n-hexadecane (31.1 ml) was charged to a bubble-tower reactorhaving 2 cm of inside diameter and 2 m of height, and 3.6 g of each ofthe above-described catalyst powders was added to make the contents ofthe reactor in a suspended state. A mixed gas of hydrogen, carbonmonoxide, and nitrogen (at a molar ratio H₂:CO:N₂ of 1:1:9) wasintroduced to pass through the bubble-tower at a flow rate of about 300ml/min. While flowing the mixed gas through the bubble-tower, thetemperature in the bubble-tower was gradually raised from roomtemperature to 220° C. within a period of several hours. At the sametime, the concentration of nitrogen in the mixed gas was graduallyreduced to a final level of zero. Then, the reaction system was held at220° C. for about 3 hours to activate the catalyst.

The reaction was conducted at a specified temperature and pressure whileintroducing the mixed gas of hydrogen, carbon monoxide, and carbondioxide at a molar ratio of H₂/CO/CO₂=47.5/47.5/5.0 and at a flow rateof 336 ml/min. (converted at a condition of normal temperature andpressure).

The obtained reaction products and non-reacted substances were analyzedby gas chromatography.

III. Reaction Conditions and Experimental Results

The reaction conditions and experimental results are shown in Tables 11and 12.

$\begin{matrix}{\begin{matrix}{{Conversion}\mspace{14mu}{of}} \\{{dimethylether}(\%)}\end{matrix} = {\frac{\begin{matrix}{{{Charge}\mspace{14mu}{rate}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} -} \\{{{Discharge}{\mspace{11mu}\;}{rate}\mspace{14mu}{of}}\mspace{11mu}} \\{\;{{non}\text{-}{reacted}\mspace{14mu}{dimethyleter}}}\end{matrix}}{{Charge}\mspace{14mu}{rate}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} \times 100}} \\{{{Yield}\mspace{14mu}{of}\mspace{14mu}{{methanol}(\%)}} = {\frac{\frac{1}{2} \times {Yielding}\mspace{14mu}{rate}\mspace{14mu}{methanol}}{{Charge}\mspace{14mu}{rate}\mspace{14mu}{of}\mspace{20mu}{dimethylether}} \times 100}} \\{{{{Yield}\mspace{14mu}{of}\mspace{14mu}{{hydrocarbons}(\%)}} = {\frac{\mspace{14mu}\begin{matrix}{\sum\lbrack {{n/2} \times {yielding}\mspace{14mu}{rate}} } \\{{of}\mspace{14mu}{hydrocarbons}} \\( {{number}\mspace{14mu}{of}\mspace{14mu}{{carbons}:n}} )\end{matrix}}{{Charge}\mspace{20mu}{rate}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} \times 100}}{{{Yield}\mspace{14mu}{of}\mspace{14mu}{CO}_{2}\mspace{14mu}(\%)} = {\frac{\frac{1}{4} \times {Yielding}\mspace{14mu}{rate}\mspace{14mu}{of}\mspace{14mu}{CO}_{2}}{{Charge}\mspace{14mu}{rate}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} \times 100}}}\end{matrix}$

TABLE 11 Example 1 Example 2 Example 3 Catalyst Catalyst No. {circlearound (1)} {circle around (2)} {circle around (3)} {circle around (1)}{circle around (4)} {circle around (1)} {circle around (4)} ParticleDensity(g/cm²) 28.2 3.17 3.20 2.82 3.31 2.82 3.31 Particle size(cm) ×10⁴ 16.9 15.6 15.5 16.9 15.2 14.4 12.9 Weight (g) 2.4 1.2 1.2 2.4 1.22.4 1.2 Value of A (g/cm) × 10⁵ 5.85 5.83 5.83 5.85 5.86 4.25 4.22Condition Temperature (° C.) 280 280 280 Pressure (kg/cm²-G) 30 30 30Reaction CO conversion (%) 40.1 53.9 52.1 Result Yield Dimethyl ether31.2 36.2 34.0 (C-mol %) Methanol 2.5 2.5 2.6 Hydrocarbons 1.2 0.8 0.5CO₂ 5.5 14.4 15.0 Dimethyl ether space time 853 990 930 yield (g/kg-cat· h)

TABLE 12 Comparative Comparative Example 1 example 1 example 2 CatalystCatalyst No. {circle around (1)} {circle around (5)} {circle around (1)}{circle around (2)} {circle around (3)} {circle around (1)} {circlearound (4)} Particle Density(g/cm²) 2.82 2.50 2.82 3.17 3.20 2.82 3.31Particle size(cm) × 10⁴ 16.9 18.4 16.9 20.1 18.5 16.9 12.9 Weight (g)2.4 1.2 2.4 2.4 1.2 2.4 1.2 Value of A (g/cm) × 10⁵ 5.85 5.85 5.85 9.688.31 5.85 4.22 Condition Temperature (° C.) 280 280 280 Pressure(kg/cm²-G) 30 30 30 Reaction CO conversion (%) 52.7 33.4 41.6 ResultYield Dimethyl ether 35.1 23.3 26.9 (C-mol %) Methanol 2.1 1.2 2.0Hydrocarbons 0.9 0.8 0.5 CO₂ 14.6 8.1 12.2 Dimethyl ether space time 960637 736 yield (g/kg-cat · h)

The catalyst for manufacturing dimethyl ether according to the presentinvention provides effects of preventing separation of individualcatalyst ingredients from each other during reaction, of assuring smoothprogress of reaction cycle, and of achieving high dimethyl ether yieldowing to the controlled particle density and particle size for each ofthe methanol synthesis catalyst, the methanol-dehydration catalyst, andthe water gas shift catalyst conforming to a derived equation.

The method for manufacturing dimethyl ether according to the presentinvention uses a slurry of solvent with a catalyst comprising a methanolsynthesis catalyst, a methanol-dehydration catalyst, and a water gasshift catalyst being integrated together, so the method provides effectsof achieving high space time yield of dimethyl ether, of being free fromproblems of plugging of catalyst and of mechanical strength of catalyst,of easiness for removing reaction heat and for controlling reactionheat, of assuring wide application range of the ratio of carbon monoxideto hydrogen, of progress of reaction under the presence of highconcentration of carbon dioxide, and of less influence of impurities andcatalyst poisons.

Embodiment 5

Synthesis of dimethyl ether proceeds following the three equilibriumreactions shown below.

CO + 2H₂

CH₃OH (1) (Methanol synthesis reaction) 2CH₃OH

CH₃OCH₃ + H₂O (2) (Dehydration reaction) H₂O + CO

H₂ + CO₂ (3) (Shift reaction)

In the case that the reaction (1) is solely carried out, the reaction iswhat is called the methanol synthesis reaction. The methanol synthesisreaction has a limitation of equilibrium, and a high pressure (80 to 120kg/cm²) is necessary to obtain a target conversion. In the single stageprocess, however, the reaction (2) which is significantly superior interms of equilibrium successively proceeds within the same reactor toconsume the methanol as the reaction product, so the inferior reactionequilibrium of reaction (1) is compensated. As a result, the dimethylether synthesis becomes very easy compared with conventional methanolsynthesis process. That is, the single stage process increases theconversion of CO/H₂.

The reaction mixture comprises non-reacted CO and H₂, reaction productsmethanol, dimethyl ether, and CO₂, and slight amount of byproducts suchas alkane. Since the composition depends on the reaction rate andequilibrium of each of the reactions (1) through (3), the single stageprocess cannot increase the amount of solely a target product. Inparticular, methanol as an intermediate is unavoidably left in thereaction products.

Reaction rate in each reaction is controlled by changing the ratio ofmethanol synthesis catalyst, dehydration catalyst, and shift catalyst.Thus the composition of the reaction products is controlled. Since,however, these three types of reactions proceed simultaneously and sinceall of these three reactions are equilibrium reactions, the control hasa limitation owing to the limit of equilibrium. With that type ofreaction system and under a normal reaction condition, it is verydifficult to attain the selectivity of dimethyl ether over 95% (in theproducts excluding CO₂).

The difficulty is described below referring to a thermodynamiccalculation.

FIG. 2 shows reaction equilibria of H₂, CO, methanol, CO₂, and water, onthe basis of reactions (1), (2), and (3). For instance, under acondition of 300° C. of reaction temperature, 50 atm of reactionpressure, and 1 of initial CO/H₂ ratio, the selectivity of dimethylether (starting material of CO, carbon molar basis, excluding CO₂) atthe reaction equilibrium is 98%. Establishing a reaction equilibrium is,however, impossible in actual process, and the selectivity of dimethylether becomes significantly lower level than the calculated value owingto the presence of methanol as an intermediate. In a state of lowertemperature than the above example, for instance at 240° C. of reactiontemperature, 50 atm of reaction pressure, and 1 of initial CO/H₂ ratio,the selectivity of dimethyl ether at a reaction equilibrium becomes 99%which is somewhat higher than that in the above example. Under thecondition, however, the rate of methanol-synthesis reaction (1) is lowand actually the reaction equilibrium is never established.

The method according to the present invention uses a mixture of methanolsynthesis catalyst, dehydration catalyst or dehydration and shiftcatalyst in the first sage reaction to yield crude dimethyl ether, anduses dehydration and/or shift catalyst in the second stage reaction toconvert most of the remained methanol into dimethyl ether, thus attainsa high selectivity of dimethyl ether.

That is, the first stage reaction increases the conversion of CO/H₂ rawmaterial mixed gas by using a combined methanol synthesiscatalyst+dehydration catalyst+shift catalyst. The second stage reactionuses dehydration catalyst and/or shift catalyst, and adopts a reactioncondition particularly suitable for the above-described reactions, thusachieving the conversion of most part of the remaining methanol intodimethyl ether to increase the selectivity of dimethyl ether.

The second stage reaction uses only dehydration and/or shift catalystwithout applying methanol synthesis catalyst. When the second stagereaction adopts a reaction condition particularly suitable for thedehydration and/or shift reaction, solely the dehydration and/or shiftreaction proceeds and only these reactions approach the equilibriumstate. In a reaction system where no methanol synthesis catalyst exists,no additional methanol yields, and the remained methanol is convertedinto dimethyl ether, so the selectivity of dimethyl ether increases.

Accordingly, the present invention provides a method for manufacturingdimethyl ether comprising the steps of: using a mixed gas as a rawmaterial gas containing carbon monoxide and either or both of hydrogenand water vapor, or further containing carbon dioxide; conducting afirst stage reaction by contacting the raw material gas with a catalystcontaining a synthesis catalyst, a methanol-dehydration catalyst, and ashift catalyst; and conducting a second stage reaction by contacting thefirst stage product gas with a catalyst containing at least one of thedehydration catalyst and the shift catalyst.

The catalyst used in the first stage according to the present inventionis a combination of known methanol synthesis catalyst, known dehydrationcatalyst, and known water gas shift catalyst. Applicable methanolsynthesis catalyst includes common industrial catalysts for methanolsynthesis, such as those of copper oxide-zinc oxide system, zincoxide-chromium oxide system, copper oxide-zinc oxide/chromium oxidesystem, copper oxide-zinc oxide/alumina system. Examples of dehydrationcatalyst are acid-base catalyst such as γ-alumina, silica,silica-alumina, and zeolite. Examples of the metallic oxide ingredientin zeolite are oxide of alkali metal such as sodium and potassium, andoxide of alkali earth metal such as calcium and magnesium. Examples ofwater gas shift catalyst are copper oxide-zinc oxide system, copperoxide-chromium oxide-zinc oxide system, and iron oxide-chromium oxidesystem. Since methanol synthesis catalyst has a strong activity as ashift catalyst, it can substitute for the water gas shift catalyst Acopper oxide supported by alumina is applicable as a dehydrationcatalyst and also as a water gas shift catalyst.

The mixing ratio of the above-described methanol synthesis catalyst,dehydration catalyst, and water gas shift catalyst is not specificallylimited, and it depends on the kind of each ingredient and reactioncondition. Normally the ratio is often in an approximate range of from0.1 to 5 wt.parts of dehydration catalyst to 1 wt.parts of methanolsynthesis catalyst, more preferably in an approximate range of from 0.2to 2; in an approximate range of from 0.2 to 5 wt.parts of water gasshift catalyst, more preferably in an approximate range of from 0.5 to3. When the methanol synthesis catalyst substitutes for the water gasshift catalyst, the above-described amount of the water gas shiftcatalyst is added to the amount of the methanol synthesis catalyst.

The catalyst used in the second stage reaction is a combination ofdehydration catalyst, dehydration-shift catalyst, and a combination ofdehydration catalyst+shift catalyst. The dehydration-shift catalyst mayuse a catalyst of copper oxide supported by γ-alumina as the catalysthaving activity of both dehydration and shifting.

Preferable catalysts used in the second stage reaction include thoselisted as examples of dehydration catalyst, shift catalyst, anddehydration-shift catalyst in the first stage reaction.

Above-described catalysts in both of first and second stage reactionsare used in a powder state. A preferable average particle size is 300 μmor less, more preferably in an approximate range of from 1 to 200 μm,and most preferably in an approximate range of from 10 to 150 μm. Toprepare powder of that range of particle size, the catalysts may furtherbe pulverized.

The type of the catalyst reactor for the first stage and the secondstage may be either a fixed bed type or a slurry bed type. When thefixed bed type reactor is applied, the catalysts are granulated to asuitable size by a known method. When the slurry bed type reactor isapplied, the applicable solvent is an arbitrary kind if only the solventmaintains liquid state under the reaction condition. Examples of thesolvent are hydrocarbons of aliphatic, aromatic, and alicyclic groups,alcohol, ether, ester, ketone, halide, or their mixture. Alternatively,gas oil after removing sulfur ingredients, vacuum gas oil, high boilingpoint distillates of coal tar after treated by hydrogenation are alsoapplicable as the solvent. The amount of catalyst in solvent depends onthe kind of solvent and the reaction condition. Normally, a preferablerange of the catalyst is from 1 to 50 wt. % to the amount of solvent,more preferably in a range of from 2 to 30 wt. %.

The mixing ratio of hydrogen and carbon monoxide may be in a range offrom 20 to 0.1 as H₂/CO molar ratio, more preferably in a range of from10 to 0.2. In the case of a mixed gas with significantly low ratio of(H₂/CO), for example, 0.1 or less, or in the case of sole carbonmonoxide without containing hydrogen, it is necessary to separatelysupply steam to conduct the shift reaction (3) in the reactor to converta part of the carbon monoxide into hydrogen and carbon dioxide. Apreferable charge rate of steam is 1 or less to the charge rate of CO. Apreferable amount of carbon dioxide yielded from the reaction is 50% orless.

A preferable condition of the first stage reaction is the reactiontemperature in a range of from 150 to 400° C., particularly in a rangeof from 200 to 350° C. The reaction temperature below 150° C. and above400° C. results in a reduction of carbon monoxide conversion. Apreferable reaction pressure is in a range of from 10 to 300 kg/cm²,particularly in a range of from 15 to 150 kg/cm². The reaction pressurebelow 10 kg/cm² results in a low conversion of carbon monoxide, and thatabove 300 kg/cm² requires special design of reactor and is uneconomicalbecause of the need of a large amount of energy for pressurizing thesystem. A preferable space velocity (charge rate of mixed gas per 1 g ofcatalyst under standard condition) is in a range of from 100 to 50000l/kg·h, and particularly preferable from 500 to 30000 l/kg·h. The spacevelocity above 50000 l/kg·h degrades the conversion of carbon monoxide,and that below 100 l/kg·h is uneconomical because of the need of anexcessively large reactor.

The condition of the second stage reaction may be the same with that ofthe first stage reaction. A major characteristic of the method accordingto the present invention is to conduct the second stage reactionseparately from the first one, which allows selecting an optimumcondition particularly for the dehydration reaction (2) and for theshift reaction (3). That is, according to an analysis of reactionequilibrium, the reactions of (2) and (3) become advantageous at a lowtemperature level, and these reactions attain sufficient levels ofreaction rate at a low temperature compared with the reaction of (1).Therefore, the second stage reaction favorably adopts lower temperaturethan the first stage reaction. Since the reaction pressure does notaffect the reaction equilibrium, low pressure level is applicable.

A practical reaction condition is in a temperature range of from roomtemperature to 300° C., more preferably in a range of from 100 to 300°C., and preferably in a pressure range of from atmospheric pressure tothe pressure of first stage reaction. Higher pressure is preferablebecause the reactor volume decreases. Nevertheless, it is not preferableto apply higher pressure than that of the first stage reaction becauseof the necessity of additional energy for pressurizing the system. Spacevelocity depends on the intensity of catalyst activity. A preferablespace velocity is normally in a range of from 100 to 50000 l/kg·h, andparticularly preferable from 500 to 30000 l/kg·h. The space velocityabove 50000 l/kg·h results in failing to fully approach to theequilibrium and failing to increase the selectivity of dimethyl ether.

The ratio of the amount of catalyst in the first and second stagesdepends on the activity of individual catalysts. A preferable ratio ofthem (first stage/second stage) is normally in a range of from 1:10 to10:1, more preferably in a range of from 1:5 to 5:1.

The CO conversion, the selectivity of dimethyl ether, and theselectivity of methanol referred to herein are defined as follows.

-   Amount of CO gas charged to reactor (Nl/min.):Fin(CO)-   Amount of CO gas discharged from reactor:Fout(CO)-   Amount of DME gas discharged from reactor:Fout(DME)-   Amount of MeOH gas discharged from reactor:Fout(MeOH)-   Amount of methane gas discharged from reactor:Fout(CH₄)=

$\begin{matrix}{( {{CO}\mspace{14mu}{conversion}} ) = \frac{{{Fin}({CO})} - {{Fout}({CO})}}{{Fin}({CO})}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} ) = \frac{2 \times {{Fout}({DME})}}{\begin{matrix}{{2 \times {{Fout}({DME})}} +} \\{{{Fout}({MeOH})} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{11mu}{methanol}} ) = \frac{{Fout}({MeOH})}{\begin{matrix}{{2 \times {{Fout}({DME})}} +} \\{{{Fout}({MeOH})} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}}\end{matrix} \times 100(\%)$

EXAMPLE Example 1

Applied apparatus is shown in FIG. 1. The apparatus comprises a seriesof a first stage reactor 1 which is a slurry bed type and a second stagereactor 2 which is a fixed bed type. To the bottom of the first stagereactor 1, a H₂ gas piping and a CO gas piping are connected, each ofwhich is equipped with an automatic pressure control valve and with amass flow rate controller (FRCA). The top of the first stage reactor 1is connected to the bottom of the second stage reactor 2. The top of thesecond stage reactor 2 is connected to a gas-liquid separator 4 via acooler 3. The gas exit of the gas-liquid separator 4 is connected to anexhaust line via a gas meter 5. The liquid exit of the gas-liquidseparator 4 is connected to a product storage tank 6. A branch islocated on the line between the exit of the reactor 2 and the cooler 3,through which the composition of exit gas is analyzed.

Catalyst A: CuO—ZnO—Al₂O₃ catalyst

Each of 185 g of copper nitrate (Cu(NO₃)₂3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃9H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about1.4 kg of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of the solutions were added dropwise to about3 liters of ion-exchanged water in a stainless steel vessel which wascontrolled at about 60° C. within about 2 hours while maintaining thecontents to pH 7.0±0.5. Then, the contents were allowed to stand forabout 1 hour for aging. When, during the treatment, the pH value wentout from a range of pH 7.0±0.5, an aqueous solution of about 1mole/liter sodium carbonate was added dropwise to keep the range of pH7.0±0.5. The resulted precipitate was filtered, and the cake was rinsedby ion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours followed bycalcining thereof in air at 350° C. for 0.5 hours to obtain the targetcatalyst. Analysis of thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=61:32:7 (by weight).

Catalyst B: CuO—Al₂O₃ catalyst

A 15.7 g of copper acetate (Cu(CH₃COO)₂H₂O) was dissolved into about 200ml of ion-exchanged water. A 95 g of γ-alumina (N612, Nikki Kagaku Co.)was further added to the mixture. The mixture was then vaporized to dry.The dried material was calcined in air at 450° C. for 4 hours. Thecalcined material was treated in hydrogen gas stream at 400° C. for 3hours to obtain a catalyst. Analysis of the catalyst gave thecomposition as Cu:Al₂O₃=5:95 (by weight).

Each of the catalyst thus prepared was pulverized in a ball mill to aparticle size of 120 μm or less.

The reactor (1) was filled with 5584 ml of n-hexadecane as the heatingmedium oil, 430 g of the catalyst A, and 215 g of catalyst B: that is,(catalyst A/catalyst B)=2/1, and (catalyst/heating medium oil)=15/100.The reactor B was filled solely with 645 g of catalyst B.

(Preliminary Reduction)

Under a condition of 10 kg/cm² of reactor pressure, 220° C. of reactortemperature, a mixed gas (H₂/N₂=1/4) was introduced to the reactor at aflow rate of 10 l/min. for 12 hours to conduct preliminary reduction.

Examples 1 through 4

A mixed gas (H₂/CO=1/1) was introduced to the reaction system at a flowrate of 80 l/min. The condition of the reactor 1 was kept unchanged at50 kg/cm² of reaction pressure and 260° C. of reaction temperature,while the condition of the reactor 2 was changed for conducting thedimethyl ether synthesis. A gas chromatograph was used for analyzing thegas composition, and a gas meter was used to determine the gas flow rateat exit of the reaction system. Thus, the CO conversion and theselectivity of each reaction product (carbon molar basis, in theproducts excluding CO₂) were computed. The result is shown in table 13.

TABLE 13 Example 1 (Comparative Example) Example 2 Example 3 Example 4Pressure in the Reactor (2) 30 10 30 reactor (2) (kg/cm²) was by-passedTemperature in the 240 240 260 reactor (2) (° C.) CO conversion(%) 34.042.4 42.2 41.4 DME selectivity(%) 67.1 96.1 94.4 93.6 Methanol 32.8 3.85.5 6.3 selectivity(%) Methane  0.1 0.1 0.1 0.1 selectivity(%)

Examples 5 through 8

A mixed gas (H₂/CO=1.5/1) was introduced to the reaction system at aflow rate of 66.7 l/min. The condition of the reactor (1) was keptunchanged at 50 kg/cm² of reaction pressure and 280° C. of reactiontemperature, while the condition of the reactor (2) was changed forconducting the dimethyl ether synthesis. With the same analytical methodas in Examples 1 through 4, the CO conversion and the selectivity ofeach reaction product were determined. The result is shown in Table 14.

TABLE 14 Example 5 (Comparative Example) Example 6 Example 7 Example 8Pressure in the Reactor (2) 30 10 30 reactor (2) (kg/cm²) was by-passedTemperature in 240 240 260 the reactor (2) (° C.) CO conversion(%) 51.566.4 64.4 64.9 DME selectivity(%) 84.4 94.9 93.8 93.2 Methanol 15.3 4.85.9 6.5 selectivity(%) Methane  0.3 0.3 0.3 0.3 selectivity(%)

The method according to the present invention yields dimethyl ether fromcarbon monoxide and hydrogen (or water vapor) at a high conversion and ahigh selectivity. As a result, dimethyl ether is obtained from thereaction products readily at a high purity, thus manufactures dimethylether at a large volume at a low cost.

Embodiment 6

The inventors studied a means to maintain a high CO conversion levelwhile decreasing the concentration of CO₂ in the recycle gas to increasethe use efficiency of raw material gas, or a means to fully remove CO₂from the non-reacted gas. A method for removing CO₂ may be the one topass the non-reacted gas containing high concentration level of CO₂through an alkali solution such as aqueous caustic soda to removethereof through reaction-absorption operation. The method is, however,not a favorable one because it needs a complex process, high investmentcost, and high operation cost to recover and reuse alkali. The inventorsfocused on the solubility and absorbability of CO2 in DME (dimethylether), and found that CO₂ is efficiently removed from the reacted gasby recycling a part of DME which was separated from CO₂ to thenon-reacted gas separation step (S2) and by using CO₂ as a materialbeing absorbed.

In concrete terms, the present invention relates to a method formanufacturing dimethyl ether which comprises: letting a raw material gascontaining at least carbon monoxide and hydrogen catalytically react toyield dimethyl ether; and letting carbon monoxide and hydrogen left inreaction products recycle for reuse thereof, wherein carbon dioxideaccompanied by the recycling carbon monoxide and hydrogen is removedtherefrom by making the recycling gas mixture contact with a liquiddimethyl ether which was separated from the reaction products to absorband remove the carbon dioxide from the recycling gas mixture.

An applicable mixing ratio of hydrogen to carbon monoxide in the rawmaterial gas is in a molar ratio (H₂/CO) ranging from 0.1 to 20, morepreferably from 0.2 to 5. Examples of the raw material gas are coalgasified gas and synthesis gas produced from natural gas.

A mixed catalyst of a methanol synthesis catalyst and amethanol-dehydration catalyst is used as the dimethyl ether-synthesiscatalyst, and, at need, further a water gas shift catalyst is added tothe mixed catalyst system. Alternatively, the water gas shift catalystmay be separately used to configure a two-stage reaction system. Thepresent invention is applicable to any of these catalysts.

An applicable methanol synthesis catalyst includes common industrialcatalysts for methanol synthesis, such as those of copper oxide-zincoxide system, zinc oxide-chromium oxide system, copper oxide-zincoxide/chromium oxide system, copper oxide-zinc oxide/alumina system.Examples of methanol-dehydration catalyst are acid-base catalyst such asγ-alumina, silica, silica-alumina, and zeolite. Examples of the metallicoxide ingredient in zeolite are oxide of alkali metal such as sodium andpotassium, and oxide of alkali earth metal such as calcium andmagnesium. Examples of water gas shift catalyst are copper oxide-zincoxide system, copper oxide-chromium oxide-zinc oxide system, and ironoxide-chromium oxide system. Since a methanol synthesis catalyst has astrong activity as a shift catalyst, it can substitute for the water gasshift catalyst. A copper oxide supported by alumina is applicable as amethanol-dehydration catalyst and also as a water gas shift catalyst.

The mixing ratio of the above-described methanol synthesis catalyst,methanol-dehydration catalyst, and water gas shift catalyst is notspecifically limited, and it depends on the kind of each ingredient andreaction condition. Normally the ratio is often in an approximate rangeof from 0.1 to 5 wt.parts of methanol-dehydration catalyst to 1 wt.partsof methanol synthesis catalyst, more preferably in an approximate rangeof from 0.2 to 2; in an approximate range of from 0.2 to 5 wt.parts ofwater gas shift catalyst, more preferably in an approximate range offrom 0.5 to 3. When the methanol synthesis catalyst substitutes for thewater gas shift catalyst, the above-described amount of the water gasshift catalyst is added to the amount of the methanol synthesiscatalyst.

The above-described catalysts are applicable to any type of reaction infixed bed, fluidized bed, and slurry-bed type reactors. In particular,the slurry-bed reaction provides uniform temperature within the reactorand yields less by-products. The following description deals with theslurry-bed reaction as a typical mode.

For the case of a slurry-bed reaction, the catalyst is used in a finepowder shape. A preferable average particle size is 300 μm or less, morepreferably in an approximate range of from 1 to 200 μm, and mostpreferably in an approximate range of from 10 to 150 μm. To preparepowder of that range of particle size, the catalysts may further bepulverized.

An applicable kind of medium oil is an arbitrary one if only the mediumoil maintains liquid state under the reaction condition. For example,the medium oil may be hydrocarbons of aliphatic, aromatic, and alicyclicgroups, alcohol, ether, ester, ketone, halide, or their mixture.Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation, Fischer-Tropsch synthesis oil, and high boiling pointfood oil are also applicable as the medium oil. The amount of catalystinsolvent depends on the kind of solvent and the reaction condition.Normally, a preferable range of the catalyst is from 1 to 50 wt. % tothe amount of solvent, more preferably in a range of from 2 to 30 wt. %.

For the reaction condition in a slurry-bed reactor, a preferablereaction temperature in the reactor is in a range of from 150 to 400°C., and particularly preferable in a range of from 250 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide. A preferable reaction pressure is in arange of from 10 to 300 kg/cm², more preferably in a range of from 15 to150 kg/cm², and most preferably in a range of from 20 to 70 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires a special design of reactorand is uneconomical because of the need of a large amount of energy forpressurizing the system. A preferable space velocity (charge rate ofmixed gas per 1 kg of catalyst under standard condition) is in a rangeof from 100 to 50000 l/kg·h, and particularly preferable from 500 to30000 l/kg·h. The space velocity above 50000 l/kg·h degrades theconversion of carbon monoxide, and that below 100 l/kg·h is uneconomicalbecause of the need of an excessively large reactor.

Thus obtained reacted gas contains carbon dioxide, carbon monoxide,hydrogen, water, and methanol, adding to dimethyl ether, and furthercontains by-products such as methane, and impurities carried in by theraw material gas. An approximate composition of the reacted gas is: 1 to40% of dimethyl ether, normally 3 to 25%; 1 to 40% of carbon dioxide,normally 3 to 25%; 10 to 70% of carbon monoxide, normally 20 to 50%; 10to 70% of hydrogen, normally 20 to 50%; 0.2 to 5% of methanol, normally0.5 to 3%; 0.05 to 0.8% of water, normally 0.1 to 0.5%; and 0 to 5% ofother ingredients.

Arbitrary means is applicable to separate dimethyl ether from thereacted gas, and a method to use difference in condensation temperatureor to use difference in boiling point is a preferable one. For the casethat the difference in condensation temperature is used, firstlymethanol and water condense during the passage of cooling of reactedgas. The condensate may be utilized for other uses or may be furtherseparated into water and methanol, and the methanol may be recycled tothe reactor. Further cooling of the reacted gas induces the condensationof dimethyl ether to which carbon dioxide dissolves thereinto, whileleaving carbon monoxide and hydrogen, which are the non-reacted gasingredients, in the gas phase. Separation of carbon dioxide fromdimethyl ether may be conducted by distillation.

According to the method of the present invention, the above-describedcarbon monoxide and hydrogen as the non-reacted gas are brought intocontact with the liquid dimethyl ether which was treated in the CO₂separator to remove the dissolved carbon dioxide. Thus the carbondioxide existing in the non-reacted gas is removed by dissolving intothe dimethyl ether. Any type of the non-reacted gas separator may beused if only the separator cools the reacted gas and dimethyl ether andmaintains the contact thereof by each other in a satisfactory degree.Examples of the separator are a shell-tube heat exchanger with a liquidholder, and a tank holding liquefied dimethyl ether with an injector ofreacted gas thereinto. By the contact of reacted gas with dimethylether, CO₂ is dissolved into and absorbed by the liquefied dimethylether. The liquefaction of dimethyl ether occurs at −25° C. underatmospheric pressure. The liquefaction becomes easy and the CO₂solubility increases with increased pressure (partial pressure ofdimethyl ether) and decreased temperature. When the separationtemperature is low, the CO₂ separation efficiency increases, but thenecessary scale of refrigerator also increases to increase theinvestment cost. Accordingly, a preferable temperature of thenon-reacted gas separator is in a range of from 0 to −70° C., morepreferably from −20 to −50° C. The pressure in separation stage maynormally be the same with that of the reaction process.

The amount of recycling dimethyl ether is in a range of from 1 to 10fold of the amount of CO₂ in the reacted gas which is brought to contactwith the liquid dimethyl ether, and more preferably in a range of from 2to 5 fold.

Based on the following definitions, the formulae 1 through 3 arederived.

-   Flow rate of CO gas fed to reactor (Nl/min.):Fin(CO)-   Flow rate of CO gas discharged from reactor:Fout(CO)-   Flow rate of DME gas discharged from reactor:Fout(DME)-   Flow rate of MeOH gas discharged from reactor:Fout(MeOH)-   Flow rate of methane gas discharged from reactor:Fout(CH)

$\begin{matrix}{({COconversion}) = {\frac{{{Fin}({CO})} - {{Fout}({CO})}}{{Fin}({CO})} \times 100(\%)}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} ) = \frac{2 \times {{Fout}({DME})}}{\begin{matrix}{{2 \times {{Fout}({DME})}} +} \\{{{Fout}({MeOH})} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{14mu}{methanol}} ) = \frac{{Fout}({MeOH})}{\begin{matrix}{{2 \times {{Fout}({DME})}} +} \\{{{Fout}({MeOH})} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}}\end{matrix}$

FIG. 3 shows an example of apparatus applied to the method according tothe present invention. The apparatus is structured while adding a DMErecycle line 111 which branches from the DME line 109 coming from theCO₂ separator S3 and which returns to the non-reacted gas separator S2.

FIG. 4 shows an enlarged view of the non-reacted gas separator S2 usedin the apparatus of FIG. 3. The non-reacted gas separator S2 comprises avessel with a cooling jacket, and a set of pipelines, namely, a reactedgas line 105 at lower portion thereof, a DME, CO₂ line 106 at bottomthereof, a DME recycle line 111 at upper portion thereof, and a recyclegas line 107 at top thereof.

EXAMPLE Example 1

A reaction experiment was conducted using an apparatus shown in FIG. 3adopting a bubble-tower reactor R (90 mm of inner diameter, 2000 mm ofheight) under a reaction condition of 280° C. and 30 kg/cm². The contentof the reactor was a slurry of 860 g of catalyst powder with 5730 g ofn-hexadecane. The unit S1 separated methanol and water from the reactorexit gas, and discharge the separated methanol and water and dischargedthereof through a methanol-water line 104. The unit S2 separated a partof the DME and CO₂ gas and discharge thereof through a DME, CO₂ line.Thus the non-reacted gas was refined. The refined non-reacted gas wasrecycled to the inlet of reactor via a recycle gas line 107. Theconsumed CO gas and hydrogen gas were supplied, and the flow ratethrough a CO gas and hydrogen gas make up gas line 101 was adjusted tomaintain the flow rate of CO gas and hydrogen gas at the inlet ofreactor to 30 Nl/min. The recycle gas contains CO₂ gas left un-separatedin the S2 unit, and the purge rate of the recycle gas was adjusted tomaintain the CO₂ gas flow rate at inlet of the reactor to 9 Nl/min. Byassuring the gas composition at the inlet of reactor to theabove-described level, the CO conversion was kept to 42%. The S2 unitwas controlled to a state of −30° C. and 30 kg/cm², and the DME fromwhich CO₂ was removed in the S3 unit was recycled at a flow rate of 0.5mole/min. through a line 111. The purge rate of the recycle gas in theabove case was 3.8%, and the loss of CO gas and hydrogen gas relative tothe make up rate was 5.0% and 3.4%, respectively.

Comparative Example 1

An experiment was conducted under the same condition with Example 1except that the DME from which CO₂ was removed in the S3 unit was notrecycled. The purge rate of the recycle gas in that case was 29.7%, andthe loss of CO gas and hydrogen gas relative to the make up rate was29.1% and 21.5%, respectively.

Example 2

Using a vessel with a cooling jacket, which is shown in FIG. 4, a testgas (CO₂:DME:CO:H₂=10:10:40:40) was introduced to the reactor at a flowrate of 50 mmole/min. through a line 103, and DME was introduced to thereactor at a flow rate of 5 mmole/min. through a line 111. At the sametime, a discharged gas was collected through a line 107, and DME waswithdrawn from the reactor through a line 106 to maintain the amount ofliquefied DME in the vessel at a fixed level. The internal pressure ofthe vessel was 50 kg/cm²G, and the internal temperature of the vesselwas −30° C. The collected gas flow rate was determined by a gas meter,and the composition thereof was analyzed by a gas-chromatograph(detector: TCD). The CO₂ in the discharged gas was 1.9 mmole/min. andthe CO₂ separation efficiency was 38%.

Example 3

An experiment was conducted under the same condition with Example 2except that the DME introduction flow rate through the line 111 was setto 10 mmole/min. The CO₂ in the discharged gas was 2.6 mmole/min. andthe CO₂ separation efficiency was 52%.

Example 4

An experiment was conducted under the same condition with Example 2except that the DME introduction flow rate through the line 111 was setto 20 mmole/min. The CO₂ in the discharged gas was 4.1 mmole/min. andthe CO₂ separation efficiency was 82%.

Comparative Example 2

An experiment was conducted under the same condition with Example 2except that the DME introduction through the line 111 was not applied.The CO₂ in the discharged gas was 1.0 mmole/min. and the CO₂ separationefficiency was 20%.

Example 5

An experiment was conducted under the same condition with Example 3except that the internal temperature of the vessel was set to −40° C.The CO₂ in the discharged gas was 3.1 mmole/min. and the CO₂ separationefficiency was 62%.

Example 6

An experiment was conducted under the same condition with Example 3except that the internal temperature of the vessel was set to −25° C.The CO₂ in the discharged gas was 2.4 mmole/min. and the CO₂ separationefficiency was 48%.

The method according to the present invention suppresses the CO₂concentration in the recycle gas to a low level, and maintains the COconversion in DME synthesis reaction to a high level. Since the methodallows to minimize the purge of the recycle gas or does not need topurge thereof, both CO and H₂ are effectively utilized. In addition, themethod allows to use the product DME as a CO₂ absorbent which is readilyseparated from the reacted products, and the process becomes a simpleone.

Embodiment 7

The present invention relates to an apparatus for manufacturing dimethylether comprising: a slurry-bed reactor filled with a dimethylether-synthesis catalyst and a medium oil therefor; a condenser forcondensing a gasified medium oil discharged from the reactor; adesulfurization tank for adsorbing to remove a catalyst-deactivationingredient from the medium oil condensed in the condenser; and a recycleline for recycling the medium oil by connecting the reactor, thecondenser, and the desulfurization tank. And relates to a method formanufacturing dimethyl ether comprising the steps of: letting a rawmaterial gas containing at least carbon monoxide and hydrogen contact aslurry of a suspended dimethyl ether-synthesis catalyst in a medium oil;cooling a reaction product gas generated from a catalytic reaction tocondense to separate a gasified medium oil carried along with thereaction product gas; obtaining dimethyl ether from the reaction productgas; removing a catalyst-deactivation ingredient from a condensed mediumoil to recycle the medium oil free of catalyst-deactivation ingredientto the slurry.

The reactor according to the present invention may be an arbitrary typeif only it is a slurry-bed type.

A dimethyl ether-synthesis catalyst according to the present inventionis a mixture of a methanol synthesis catalyst and a methanol-dehydrationcatalyst, and, if needed, a water gas shift catalyst is further added tothe catalyst system. These catalyst ingredients may be used in a mixedstate, or the water gas shift catalyst may be separated from other twocatalyst ingredients to configure a two stage reaction system. Thepresent invention is applicable for any types of the above-describedcatalyst systems.

Examples of a methanol synthesis catalyst are copper oxide-zinc oxidesystem, zinc oxide-chromium oxide system, copper oxide-zincoxide/chromium oxide system, copper oxide-zinc oxide/alumina systemwhich catalyst systems are commonly used in industrial methanolsynthesis. Examples of a methanol-dehydration catalyst are acid-basecatalyst such as γ-alumina, silica, silica-alumina, and zeolite.Examples of the metallic oxide ingredient in zeolite are oxide of alkalimetal such as sodium and potassium, and oxide of alkali earth metal suchas calcium and magnesium. Examples of a water gas shift catalyst arecopper oxide-zinc oxide system, copper oxide-chromium oxide-zinc oxidesystem, and iron oxide-chromium oxide system. Since a methanol synthesiscatalyst has a strong activity as a shift catalyst, it can substitutefor the water gas shift catalyst. A copper oxide supported by alumina isapplicable as dehydration catalyst and also as water gas shift catalyst.

The mixing ratio of above-described methanol synthesis catalyst,methanol-dehydration catalyst, and water gas shift catalyst is notspecifically limited, and it depends on the kind of each ingredient andreaction condition. Normally the ratio is often in an approximate rangeof from 0.1 to 5 wt.parts of methanol-dehydration catalyst to 1 wt.partsof methanol synthesis catalyst, more preferably in an approximate rangeof from 0.2 to 2; in an approximate range of from 0.2 to 5 wt.parts ofwater gas shift catalyst, more preferably in an approximate range offrom 0.5 to 3. When the methanol synthesis catalyst substitutes for thewater gas shift catalyst, the above-described amount of the water gasshift catalyst is added to the amount of the methanol synthesiscatalyst.

The above-described catalysts are used in a powder state. A preferableaverage particle size is 300 μm or less, more preferably in anapproximate range of from 1 to 200 μm, and most preferably in anapproximate range of from 10 to 150 μm. To prepare powder of that rangeof particle size, the catalysts may further be pulverized.

An applicable kind of medium oil is arbitrary if only the medium oilmaintains liquid state under the reaction condition. Examples of themedium oil are hydrocarbons of aliphatic, aromatic, and alicyclicgroups, alcohol, ether, ester, ketone, halide, or their mixture.Alternatively, gas oil after removing sulfur ingredients, vacuum gasoil, high boiling point distillates of coal tar after treated byhydrogenation, Fischer-Tropsch synthesis oil, and high boiling pointfood oil are also applicable as the medium oil. The amount of catalystin the solvent depends on the kind of solvent and the reactionconditions. Normally, a preferable range of the catalyst is from 1 to 50wt. % to the amount of solvent, more preferably in a range of from 2 to30 wt. %.

The condenser according to the present invention is arbitrary if only itcondenses vaporized medium oil. A heat exchanger or other types may beapplied.

Removal of catalyst deactivation ingredients is preferably carried outusing an adsorbent An applicable type of the adsorbent may be thecommonly used one such as γ-alumina, activated carbon, and zeoliteadsorber. A necessary amount of the adsorbent depends on the adsorptioncapacity of the applied adsorbent. The number of adsorbers is preferablymore than one to ensure continuous adsorption operation duringregeneration cycle of an adsorber.

The recycle line establishes a recycle passage through the slurry-bedreactor, the condenser, and the adsorber. A means to prevent backflow islocated in the line. The backflow-preventive means may be a simple checkvalve or may be a forced circulation system using a pump.

The apparatus according to the present invention is further providedwith a storage tank, an intermediate tank, valves, and instruments, asneeded.

The molar mixing ratio of hydrogen and carbon monoxide, (H₂/CO), may bein a range of from 0.5 to 3.0, more preferably from 0.8 to 2.0. In thecase of a mixed gas with significantly low ratio of (H₂/CO), forexample, 0.5 or less, or in the case of sole carbon monoxide withoutcontaining hydrogen, it is necessary to separately supply steam toconvert a part of the carbon monoxide into hydrogen and carbon dioxidewithin the reactor. The amount of steam supply is equal to the amount ofcarbon monoxide to be converted (equal to the insufficient amount ofhydrogen). The amount of carbon dioxide becomes the same molar amount ofthe converted carbon monoxide. Examples of that type of raw material gasare coal gasified gas, synthesis gas derived from natural gas, andmethane in coal stratum. When sulfur compounds exist in the raw materialgas, a preliminary desulfurization is needed to prevent catalystdeactivation. The desulfurization treatment reduces the concentration ofsulfur compounds to several hundreds of ppm or less, normally in anapproximate range of from 50 to 200 ppm. Sulfur compounds include SOx,H₂S, and COS.

A preferable reaction temperature in the slurry-bed reactor is in arange of from 150 to 400° C., and particularly preferable in a range offrom 250 to 350° C. The reaction temperature below 150° C. and above400° C. degrades the conversion of carbon monoxide. A preferablereaction pressure is in a range of from 10 to 300 kg/cm², morepreferably in a range of from 15 to 150 kg/cm², and most preferably in arange of from 20 to 70 kg/cm². The reaction pressure below 10 kg/cm²results in a low conversion of carbon monoxide, and that above 300kg/cm² requires a special design of reactor and is uneconomical becauseof the need of a large amount of energy for pressurizing the system. Apreferable space velocity (charge rate of mixed gas per 1 kg of catalystunder standard condition) is in a range of from 100 to 50000 l/kg·h, andparticularly preferable from 500 to 30000 l/kg·h. The space velocityabove 50000 l/kg·h degrades the conversion of carbon monoxide, and thatbelow 100 l/kg·h is uneconomical because of the need of an excessivelylarge reactor.

A condition for eliminating the catalyst deactivation ingredients is toreduce the concentration of catalyst deactivation ingredients remainingin a medium oil after removal thereof by adsorption to 100 ppm or less,and more preferably to 50 ppm or less.

EXAMPLE Example 1

FIG. 5 shows an apparatus for manufacturing dimethyl ether as an exampleaccording to the present invention. The apparatus comprises a slurryreactor 201, a condenser 202, a gas-liquid separator 203, an adsorber204, and a pump 205. These units are connected each other by pipes toform a recirculation line 206. A raw material gas is charged to thebottom of the slurry reactor 201 via a raw material gas feed pipe 207. Amedium oil which is vaporized from the slurry reactor 201 or which isdischarged from the slurry reactor 201 along with a high temperaturereacted gas is cooled to a temperature of condensing point or below byexchanging heat thereof in the condenser 202. The cooled and partiallycondensed medium oil enters the gas-liquid separator 203 to separatethereof to reacted products and non-reacted gas. Thus separated andcollected medium oil is sent to the adsorber 204, where the catalystdeactivation ingredients are adsorbed to remove thereof from the mediumoil. The clean medium oil is then recycled to the reactor 201 by thepump 205 to maintain the concentration of catalyst deactivationingredients in the medium oil at a low level. The desulfurization in theadsorption-desulfurization tank 204 may be carried out under the samehigh pressure condition with that in the reactor or may be conducted atatmospheric pressure after depressurizing. The number of adsorbers 204may be two or more for alternative use.

An gas-flow type autoclave was used as the slurry-bed reactor. Thereaction experiment was carried out under the conditions of 30 kg/cm²Gof pressure, 280° C. of temperature. The charged raw material gascontained carbon monoxide and hydrogen at a ratio of 1:1. The ratio ofthe weight of catalyst to the molar flow rate of carbon monoxide wasW/F=4.0 g-cat.·h/CO-mole. The used medium oil contained different levelsof dissolved sulfur ingredient concentration for each run. The appliedcatalyst was a mixture of methanol synthesis catalyst ofcopper-zinc-alumina system and methanol-dehydration catalyst ofcopper-alumina system at a weight ratio of 2 to 1, both in powder form.

A medium oil was n-cetane which contained no sulfur ingredient. A mediumoil was industrial gas oil 1 which was refined to a level of 250 ppm ofsulfur ingredients. A medium oil was industrial gas oil 2 which wasrefined to a level of 100 ppm of sulfur ingredients. A medium oil wasindustrial gas oil 3 which was refined to a level of 50 ppm of sulfuringredients.

The resulted conversion of carbon monoxide was: 50.5% in n-cetane mediumoil, 8.1% in industrial gas oil 1 containing 250 ppm of sulfuringredients, 35.2% in industrial gas oil 2 containing 100 ppm of sulfuringredients, and 45.6% in industrial gas oil 3 containing 50 ppm ofsulfur ingredients. Therefore, according to the present invention, theconcentration of sulfur ingredients in a medium oil after the sulfuringredients were removed by adsorption is necessary to 100 ppm or less,and preferably to 50 ppm or less.

The above-described experimental result is shown in Table 15. Thereaction ratio is expressed by the conversion of carbon monoxide, a rawmaterial, on molar basis.

TABLE 15 Concentration of Carbon monoxide dissolved sulfur conversionMedium oil ingredients (ppm) (C-mol %) Industrial gas oil 1 250 8.1Industrial gas oil 2 100 35.2 Industrial gas oil 3 50 45.6

For the case that a n-cetane having 286.8° C. of boiling point and 226of molecular weight was used as the medium oil, and that the internalcondition of the reactor was set to 30 kg/cm²G, 280° C., W/F (ratio ofcatalyst weight to raw material gas carbon monoxide)=6 kg-cat/CO-kg.mol,and 20 wt. % of slurry concentration, the calculated amount of mediumoil discharged from the reactor determined from a vapor-liquidequilibrium relation under the internal condition of reactor results inabout 50%/h in terms of weight ratio to the initially existing amount ofthe medium oil in the reactor.

The actual discharge amount of the medium oil is, however, less than thecalculated value owing to the temperature decrease at upper part of thereactor and to other variables. Nevertheless, under a normal condition,sufficient amount of medium oil for rectifying thereof recirculates tothe adsorber 204 in FIG. 5.

Example 2

Same procedure as in Example 1 was conducted using the medium oilslisted in table 16. The result is shown in Table 16.

TABLE 16 Conversion Concentration of dissolved of carbon monoxide Mediumoil metallic carbonyl (ppm) (C-mol %) Industrial gas oil 4 230 (ironcarbonyl) 9.2 Industrial gas oil 5  51 (iron carbonyl) 32.7 Industrialgas oil 6  3 (iron carbonyl) 44.8 Industrial gas oil 7 190 (nickelcarbonyl) 7.6 Industrial gas oil 8  45 (nickel carbonyl) 31.1 Industrialgas oil 9  2 (nickel carbonyl) 43.6

Example 3

A continuous flow high pressure reaction gas was used to suppresscontinuously the concentration of sulfur ingredients in the medium oil.The effectiveness of the apparatus and method according to the presentinvention was confirmed.

The applied apparatus is shown in FIG. 6. Adding to the configuration ofthe apparatus in FIG. 5, a second condenser 208 and a second gas-liquidseparator 209 are located to the gas exhaust line of the gas-liquidseparator 203.

The reacted and refined gas which was separated from medium oil in thegas-liquid separator 203 is cooled in the second condenser 208, and isthen separated into liquid phase containing dimethyl ether, methanol,and water, and gas phase containing non-reacted gas and carbon dioxidein the second gas-liquid separator 209. The liquid phase is recoveredthrough a line 210, and the gas phase is recovered through a line 211.

The reaction condition was the same with that in Example 1. The rawmaterial charged to the reactor was a mixture of carbon monoxide andhydrogen at 1:1 ratio containing 600 ppm of hydrogen sulfide as theimpurity. The applied medium oil was n-cetane containing no sulfuringredient. The adsorbent used was granular γ-alumina. A sufficientvolume of adsorbent to adsorb and desulfurize the hydrogen sulfidedissolved in the medium oil was charged into the adsorber 204.

The result of the reaction is shown in FIGS. 7 and 8.

FIG. 7 shows the change in concentration of sulfur ingredients in themedium oil for the case of empty adsorber 204 and of adsorber 204containing adsorbent. For the case of application of desulfurizationtreatment, the sulfur concentration in the recycling medium oil is keptto near zero level. For the case of non-application of desulfurizationtreatment, however, the sulfur concentration in the medium oil increaseswith the operating time.

FIG. 8 shows the change in conversion of carbon monoxide as a rawmaterial responding to the change in sulfur ingredients with time shownin FIG. 7. For the case of application of desulfurization treatment, theconversion of carbon monoxide as a raw material keeps nearly the initiallevel. For the case of non-application of desulfurization treatment,however, the conversion of carbon monoxide as a raw material shows anabrupt decrease with time. The phenomena prove the effectiveness of thepresent invention.

According to the apparatus and method of present invention, a rawmaterial mixed gas consisting mainly of carbon monoxide in a coalgasified gas, a synthesis gas produced from natural gas, or the like,and hydrogen is charged to a slurry of reaction catalyst with medium oilin a reactor to synthesize dimethyl ether, wherein the slight amount ofcatalyst deactivation ingredients existing in the raw material gas iscontinuously removed to maintain the concentration of catalystdeactivation ingredients in the medium oil to a low level, thus keepinga high catalyst activity for a long period.

Embodiment 8

The present invention relates to a slurry-bed reactor which comprises: acatalyst reaction layer in a state of slurry of a suspending powdercatalyst with a high boiling point medium oil, charging a reaction rawmaterial gas thereto at a bottom portion of the reactor; a condenser tocondense the medium oil vaporized from the catalyst slurry layer insideof the reactor, wherein the condenser and the catalyst slurry layer areseparated from each other by a separation member having a liquid sealmechanism to form a condensation chamber.

The reactor according to the present invention is a slurry-bed reactor,and the section for holding catalyst slurry as the slurry-bed isequipped with a cooling mechanism for removing reaction heat.

At an upper space of the reactor, a condenser is installed to condensethe high boiling point medium oil vaporized from the catalyst slurrybed. The condenser may be an arbitrary type, and a preferable one is amulti-tube type. When the upper space in the reactor is insufficient forinstalling the condenser, the upper portion of the reactor may beextended in length or extending in diameter to secure the space for theinstallation area thereof.

The condenser and the catalyst slurry bed are separated from each otherby a separation member having a liquid seal mechanism to create acondensation chamber that is separated from the catalyst slurry bed. Theliquid seal mechanism means a mechanism that allows passing a gas freelytherethrough in an empty state and that prohibits gas passing freelytherethrough when the mechanism is filled with liquid while the gas isallowed to pass the filled liquid by a pressure difference or bydissolving into and vaporizing from the filled liquid. An example of apreferred liquid seal member is a bubble-cap plate which is used in adistillation column.

There is a need to prepare a return line for smoothly returning themedium oil condensed in the condensation chamber to the catalyst slurrylayer. The return line is required to maintain the independence of thecondensation chamber from the catalyst slurry layer. Accordingly, theline has a sealed structure separate from the space between the catalystslurry layer and the condensation chamber. In addition, for assuring theindependence of the line exit, it is preferable to apply a liquid sealmechanism. In that case, the discharge opening of the liquid sealmechanism is necessary at a lower level than the liquid level in theliquid seal mechanism of the separation member so as the condensedliquid to be collected to the return line and to be discharged into thecatalyst slurry layer. It is also preferable that the lower end of thereturn line is positioned inside of the catalyst slurry layer to let thecatalyst slurry layer function as the liquid seal mechanism. In any way,it is preferable that the return of condensed liquid to the catalystslurry layer is done by a natural flow down motion.

The sections in the reactor other than the above-described sections maybe similar with those in conventional reactors. The raw material gascharge line is connected to the bottom or near the bottom of thereactor. The reaction product gas discharge line is connected to the topor near the top of the reactor. For the case of dimethyl ether-synthesisreactor, the recycle line of a mixed gas of carbon monoxide and hydrogenseparated from the reaction product gas is connected to either thepassage of the raw material gas charge line or the reactor directly. Inaddition, instruments such as a pressure gauge and a thermometer, and ifnecessary, an agitator and an auxiliary raw material charge line may bemounted to the reactor.

A dimethyl ether-synthesis apparatus that contains the dimethylether-synthesis reactor using the reactor according to the presentinvention may be the same configuration with conventional apparatusexcept that the condenser to condense the medium oil and the gas-liquidseparator to separate the condensed medium oil are not required. Thatis, at the exit of the reactor, a series of equipment are connected in asequent order of: a methanol-water separator which cools the reactionproduct gas to condense methanol and water to separate them from the gasphase; a non-reacted gas separator which further cools the gas tocondense dimethyl ether and carbon dioxide, thus separating thereof fromcarbon monoxide and hydrogen; and a CO₂ separator which separatesdimethyl ether and carbon dioxide from the condensed mixture. Each ofthe methanol-water separator and the non-reacted gas separator mayfurther be divided into a condenser and a gas-liquid separator.Alternatively, methanol, water, dimethyl ether, and carbon dioxide maybe condensed or solidified together, then carbon monoxide and hydrogenmay be separated, followed by separating dimethyl ether from thecondensed solid.

FIG. 11 shows an example of the apparatus implementing theabove-described alternative method. The apparatus comprises a reactor R,a methanol-water separator S1, a non-reacted gas separator S2, and a CO₂separator S3. A raw material gas charge line 302 is connected to thebottom of the reactor R To the raw material gas charge line 302, amake-up (fresh) gas line 301 through which the fresh raw material gas issupplied, and a recycle gas line 307 through which the non-reacted COand H₂ gases are supplied are connected. A reaction product gas line 303to discharge the reaction products connects the top of the reactor R tothe inlet of the methanol-water separator S1. A methanol-water line 304is connected to the exit of methanol-water separator S1. A reactionproduct gas line 305 is connected to the exit of reaction products onthe methanol-water separator S1. The other end of the reaction productgas line 305 is connected to the inlet of the non-reacted gas separatorS2. The other end of the recycle gas line 7 is connected to the exit ofnon-reacted gas on the non-reacted gas separator S2. The recycle gasline 7 is provided with a branched purge line 310 to draw out a part ofthe gas. A DME, CO₂ line 306 is connected to the exit of DME, CO₂, onthe non-reacted gas separator S2. The other end of the DME, CO₂ line 306is connected to the CO₂ separator S3. A CO₂ line 308 is connected to theCO₂ exit on the CO₂ separator S3. A DME line 309 is connected to theexit of DME on the CO₂ separator S3.

When the reactor according to the present invention is used as thedimethyl ether-synthesis reactor, a mixed catalyst of a methanolsynthesis catalyst and a methanol-dehydration catalyst is used as thedimethyl ether-synthesis catalyst to fill the reactor, and, if needed, awater gas shift catalyst is further added to the mixed catalyst system.

An applicable kind of medium oil is arbitrary if only the medium oilmaintains liquid state under the reaction condition. In the case ofdimethyl ether synthesis, for example, the medium oil may behydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture. Alternatively, gas oilafter removing sulfur ingredients, vacuum gas oil, high boiling pointdistillates of coal tar after treated by hydrogenation, Fischer-Tropschsynthesis oil, and high boiling point food oil are also applicable asthe medium oil. The amount of catalyst in the solvent depends on thekind of solvent and the reaction condition. Normally, a preferable rangeof the catalyst is from 1 to 50 wt. % to the amount of solvent, morepreferably in a range of from 2 to 30 wt. %.

For the case of a dimethyl ether synthesis reaction, a preferablereaction temperature in the reactor is in a range of from 150 to 400°C., and particularly preferable in a range of from 250 to 350° C. Thereaction temperature below 150° C. and above 400° C. degrades theconversion of carbon monoxide. A preferable reaction pressure is in arange of from 10 to 300 kg/cm², more preferably in a range of from 15 to150 kg/cm², and most preferably in a range of from 20 to 70 kg/cm². Thereaction pressure below 10 kg/cm² results in a low conversion of carbonmonoxide, and that above 300 kg/cm² requires special design of reactorand is uneconomical because of the need of large amount of energy forpressurizing the system. A preferable space velocity (charge rate ofmixed gas per 1 kg of catalyst under standard condition) is in a rangeof from 100 to 50000 l/kg·h, and particularly preferable from 500 to30000 l/kg·h. The space velocity above 50000 l/kg·h degrades theconversion of carbon monoxide, and that below 100 l/kg·h is uneconomicalbecause of the need of an excessively large reactor.

EXAMPLE

FIG. 9 shows the structure of reactor as an example according to thepresent invention. FIG. 10 shows a partial-enlarged view of the upperportion of the reactor.

The body 313 of the reactor R is in a cylindrical shape. A slurry-bedchamber 311 located at the lower part of the body 313 is filled with acatalyst slurry 318. A heat transfer tube 317 is located in the catalystslurry phase 318 to remove reaction heat. The reactor body 313 isextended upward, and a condensation chamber 312 at the upper portion ofthe body 313 contains a multi-tube condenser 314 to condense a mediumoil, to which condenser 314 a coolant pipe 319 is connected. Thecondensation chamber 312 and the slurry-bed chamber 311 are separatedfrom each other by a bubble-cap plate 315. The bubble-cap plate 315 hasa structure of a plate with multiple short-pipes 320 penetratingtherethrough, with a cap 321 on each of the short pipes 320. In anoperating state, the medium oil condensed in the condenser 314 fallsdropwise onto the plate to be held thereon, which hold-up fills the gapbetween the short pipe 320 and the cap 321 to function as a liquid seal.A downcomer 316 is located at one side of the plate 315 to return thecondensed medium oil to the catalyst slurry layer 318, and the lower endof the downcomer enters into the catalyst slurry layer 318. A rawmaterial gas line 302 is connected to the bottom of the reactor body313, and a reaction product gas line 303 is connected to the topthereof. A mixed powder of a methanol synthesis catalyst and amethanol-dehydration catalyst is used in the reactor R, and further, atneed, powder of a water gas shift catalyst is mixed thereinto. To thecatalyst, a high boiling point medium oil such as n-cetane having 286.8°C. of boiling point and 226 of molecular weight is added to form aslurry of the catalyst with liquid. Concentration of the slurry is ataround 20 wt. %. A raw material gas such as synthesis gas containingcarbon monoxide and hydrogen as main raw materials is supplied throughthe pipe 302 into the lower part of the reactor R. The raw material gasreacts during the ascending passage through the catalyst slurry layer318, and the reacted products are discharged from the pipe 303 at top ofthe reactor R. The reaction is carried out under a condition of 30kg/cm²G, 280° C., and W/F=6 kg.cat/CO.kg.mole of raw material gas chargerate (W/F is the ratio of the weight of catalyst to the charge rate ofraw material carbon monoxide). Since the dimethyl ether synthesis is asignificantly exothermic reaction, a heat transfer tube 317 is locatedin the catalyst slurry layer 318 within the reactor R to remove thereaction heat. During the operation of the reactor, the upper side ofthe bubble-cap plate 315 is covered with the cooled and condensed highboiling point medium oil. As a result, a part of the vapor of highboiling point medium oil came from lower section of the reactor iscaught by the liquid high boiling point medium oil during the passage ofcrossing the bubble-cap plate 315. The vapor of high boiling pointmedium oil which was not caught by the liquid on the bubble-cap plate315 is cooled in the condenser 314 to fall onto the bubble-cap plate 315as droplets. The high boiling point medium oil thus efficiently cooledand separated from the reaction product gas returns to the catalystslurry layer passing through a downcomer 316 which connects thebubble-cap plate 315 with the catalyst slurry layer 318 and whichimmerses the bottom end thereof into the catalyst slurry layer 318. Alow temperature heating medium may be introduced to the condenser 314 tocool the ascended vapor of high boiling point medium oil, or a lowtemperature process fluid such as the raw material gas may be used asthe cooling medium.

If no cooling for recovering the medium oil is applied, the n-cetaneconcentration in the discharged gas reaches to as high as 2 to 4 mole %.Since the volume of discharged gas is large, the net amount ofdischarged n-cetane becomes large.

According to the present invention, there is no need to recycle thedischarged medium oil to the reactor, and the operation of the reactoris stabilized. In particular, the structure of reaction product gas lineof the reactor becomes significantly simple, and there is no need toapply high pressure pump nor gas-liquid separator for recycling themedium oil.

Embodiment 9

Regarding the dimethyl ether-synthesis reaction in a slurry-bed reactor,six molar volumes of raw material gas changes to two molar volumes ofreaction product gas when the reaction proceeded to 100%. Since,however, actual one-pass conversion of the raw material gas is at about50%, six molar volumes of raw material gas changes to one fold of volumeof reaction product gas, and three fold of volume of reaction gas isdischarged from the reactor as the non-reacted gas. That is, the gasvolume reduces to one third.

According to the present invention, the principle of the above-describedgas volume reduction is utilized. In concrete terms, the internal of thereactor is divided into two sections: upper reaction tank and lowerreaction tank. The lower tank firstly makes the non-reacted recycle gasreact, thus reducing the volume of non-reacted gas flowing through thereactor. Then, the gas is added to the make up gas to let these gasesreact in the upper tank.

Therefore, according to the present invention, a reactor with a smallerinner diameter maintains the flow condition in the reactor at a uniformbubble-tower state owing to the reduction of the gas volume by reactingthe non-reacted recycle gas firstly in the reactor.

The reactor according to the present invention is a slurry-bed reactorwhich is divided into two tanks: upper tank and lower tank Each of theupper tank and the lower tank has a slurry-bed holding section and anupper space. The slurry-bed holding section, or the catalyst slurryholding section, is provided with a cooling mechanism to remove reactionheat. The boundary of the upper tank and the lower tank is necessary tohave a structure that the slurry in the upper tank does not flow downinto the lower tank and that the reaction product gas in the lower tankenters into the slurry-bed in the upper tank. A structure having thenecessary functions is to open a connection hole on the bottom plate ofthe upper tank while mounting a check valve to the connection hole toprevent the slurry in the upper tank from flowing down into the lowertank. An alternative structure is to install a connection pipe betweenthe upper space of the lower tank and the upper tank, which connectionpipe is extended to a height that the slurry in the upper tank does notflow down into the lower tank even when the reaction stops. The latterstructure may further be assured its function by installing anopen/close valve to the connection pipe. The connection pipe may beinstalled either inside or outside of the tanks. For both structures,there is no need for the separation wall between the upper tank and thelower tank to have pressure-resistant performance as a high pressurevessel because the pressure difference between these tanks is withinseveral atm.

The raw material gas feed pipe through which the fresh raw material gasis charged to the reactor may be connected to either of the upper spaceof the lower tank or the slurry-bed of the upper tank.

Other sections than those described above may be similar with those inconventional reactors, and instruments such as a pressure gauge and athermometer, and if necessary, an agitator and an auxiliary raw materialcharge line may be mounted to the reactor.

A dimethyl ether-synthesis apparatus that contains the dimethylether-synthesis reactor using a reactor according to the presentinvention may be the same configuration with conventional apparatus.That is, at the exit of the reactor, a series of equipment are connectedin a sequent order of: a condenser to condense the medium oil vaporizedfrom the reactor; a gas-liquid separator to separate the condensedmedium oil; a methanol-water separator which cools the reaction productgas to condense methanol and water to separate them from the gas phase;a non-reacted gas separator which further cools the gas to condensedimethyl ether and carbon dioxide, thus separating thereof from carbonmonoxide and hydrogen; and a CO₂ separator which separates dimethylether and carbon dioxide from the condensed mixture. Each of themethanol-water separator and the non-reacted gas separator may furtherbe divided into a condenser and a gas-liquid separator. Alternatively,methanol, water, dimethyl ether, and carbon dioxide may be condensed orsolidified together, then carbon monoxide and hydrogen may be separated,followed by separating dimethyl ether from the condensed solid.

FIG. 14 shows an example of the apparatus implementing theabove-described method. The apparatus comprises a reactor R, amethanol-water separator S1, a non-reacted gas separator S2, and a CO₂separator S3. A make up (fresh) gas line 401 is connected to the centerpart of the reactor R. A recycle gas line 402 through which thenon-reacted CO and H2 are recycled to charge is connected to the bottomof the reactor. A reaction product gas line 403 to discharge thereaction products connects the top of the reactor R to the inlet of themethanol-water separator S1, (the condenser for medium oil and thegas-liquid separator therefor are installed upstream of themethanol-water separator S1, if needed). A methanol-water line 404 isconnected to the exit of methanol-water on the methanol-water separatorS1. A reaction product gas line 405 is connected to the exit of reactionproducts on the methanol-water separator S1. The other end of thereaction product gas line 405 is connected to the inlet of thenon-reacted gas separator S2. The other end of the recycle gas line 402is connected to the exit of non-reacted gas on the non-reacted gasseparator S2. The recycle gas line 402 is provided with a branched purgeline 407 to withdraw a part of the gas. A DME, CO₂, line 406 isconnected to the exit of DME, CO₂ on the non-reacted gas separator S2.Other end of the DME, CO₂, line 406 is connected to the CO₂ separatorS3. A CO₂ line 408 is connected to the CO₂ exit on the CO₂ separator S3.A DME line 409 is connected to the exit of DME on the CO₂ separator S3.

A mixed catalyst of a methanol synthesis catalyst and amethanol-dehydration catalyst is used as the dimethyl ether-synthesiscatalyst to fill the reactor according to the present invention, and, ifneeded, a water gas shift catalyst is added to the mixed catalystsystem.

An applicable kind of medium oil is arbitrary if only the medium oilmaintains liquid state under the reaction condition. In the case ofdimethyl ether synthesis, for example, the medium oil may behydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture. Alternatively, gas oilafter removing sulfur ingredients, vacuum gas oil, high boiling pointdistillates of coal tar after treated by hydrogenation, Fischer-Tropschsynthesis oil, and high boiling point food oil are also applicable asthe medium oil. The amount of catalyst in the solvent depends on thekind of solvent and the reaction condition. Normally, a preferable rangeof the catalyst is from 1 to 50 wt. % to the amount of solvent, morepreferably in a range of from 2 to 30 wt. %.

For the case of dimethyl ether synthesis reaction, a preferable reactiontemperature in the reactor is in a range of from 150 to 400° C., andparticularly preferable in a range of from 250 to 350° C. The reactiontemperature below 150° C. and above 400° C. degrades the conversion ofcarbon monoxide. A preferable reaction pressure is in a range of from 10to 300 kg/cm², more preferably in a range of from 15 to 150 kg/cm², andmost preferably in arrange of from 20 to 70 kg/cm². The reactionpressure below 10 kg/cm² results in a low conversion of carbon monoxide,and that above 300 kg/cm² requires special design of reactor and isuneconomical because of the need of large amount of energy forpressurizing the system. A preferable space velocity (charge rate ofmixed gas per 1 kg of catalyst under standard condition) is in a rangeof from 100 to 50000 l/kg·h, and particularly preferable from 500 to30000 l/kg·h. The space velocity above 50000 l/kg·h degrades theconversion of carbon monoxide, and that below 100 l/kg·h is uneconomicalbecause of the need of an excessively large reactor.

EXAMPLE

FIG. 12 shows the structure of reactor as an example according to thepresent invention.

A reactor R is totally in a longitudinal cylindrical shape. Inside ofthe reactor is divided into an upper tank 411 and a lower tank 412 witha dish-shaped separation plate 413. At the center of the separationplate 413, a connection pipe 414 having a check valve 415 is located toconnect the upper tank with the lower tank while preventing theslurry-bed of the upper tank from flowing down. A gas distributor 416 islocated at the upper end of the connection pipe 414 in the upper tank.Both of the upper tank 411 and the lower tank 412 are filled withslurry-bed. Each of the slurry-beds 417, 418 has a heat transfer tube419, 420, respectively, to remove the reaction heat.

According to the reactor, only the non-reacted gas is pumped to thebottom of the lower tank 412 of the reactor through a recycle gas line402. The raw material make up gas is supplied through a make up gas line401 into the upper space of the catalyst slurry layer 418 in the lowertank 412 of the reactor to mix with the reaction product gas coming fromthe non-reacted recycle gas line. The mixed gas then passes through theconnection pipe 414 in the reactor, enters the upper tank 411 of thereactor, and after reacted, flows out from the reaction product gas line403.

When the reactor treats 1000 Nm³/min. of make up raw material gas andfour fold of volume of non-reacted recycle gas at an one-pass conversionof 50% under a condition of 30 kg/cm²G of reaction pressure and 300° C.of reaction temperature, the non-reacted recycle gas of 4000 Nm³/min.reduces its volume in the lower tank 412 of the reactor to 4000×2/3=2667Nm³/min. As a result, the gas is mixed with the make up raw material gasof 1000 Nm³/min. to become a volume of 3667 Nm³/min. That is, the uppertank 411 to which the maximum volume of gas enters in the reactoraccepts the gas of 3667 Nm³/min.

Under a sustained state of internal flow condition of the reactor, thecross sectional area of the reactor is proportional to the maximum gasflow rate. Since the cross sectional area of the reactor is proportionalthe square of the inner diameter thereof, the inner diameter of thereactor is proportional to the square root of the maximum gas flow rate.Accordingly, the example calculation given above results in about 10.4 mof inner diameter of reactor to treat the gas of 5000 Nm³/min. comparedwith about 12.2 m of inner diameter in conventional reactor for the samegas flow rate. The difference is a significant merit for fabricating acommercial large scale reactor.

FIG. 13 is another example of reactor according to the presentinvention. For the reactor R, a connection pipe 414 comes out from theupper space of the lower tank to outside of the reactor, and extendsupward to the upper level than the surface of the slurry-bed 417 in theupper tank 411, then connects with a make up gas line 401 to enter theupper tank 411 down into the bottom section of the upper tank 411, andfinally joins to a gas distributor 416. The other structure is the samewith that of FIG. 12.

A reactor according to the present invention synthesizes dimethyl etherby charging a raw material gas such as coal gasified gas and synthesisgas produced from natural gas, containing carbon monoxide and hydrogenas the main raw materials, to a slurry reactor in which a reactioncatalyst is suspended in a medium oil. By decreasing the maximum gasflow rate inside of the reactor relative to the same throughput, thereactor size is decreased, thus reducing the fabrication cost of thehigh pressure reactor, which cost occupies a large portion of theinvestment cost in a high pressure process.

Embodiment 10

Synthesis of dimethyl ether proceeds following the three equilibriumreactions shown below.

(1) (Methanol synthesis reaction)

(2) (Dehydration reaction)

(3) (Shift reaction)

In the case that the reaction (1) is solely carried out, the reaction iswhat is called the methanol synthesis reaction. The methanol synthesisreaction has a limitation of equilibrium, and a high pressure (80 to 120kg/cm²) is necessary to obtain a target conversion. In the single stageprocess, however, the reaction (2) which is significantly superior interms of equilibrium successively proceeds within the same reactor toconsume the methanol as the reaction product, so the inferior reactionequilibrium of reaction (1) is compensated. As a result, the dimethylether synthesis becomes very easy compared with conventional methanolsynthesis process. In other words, the single stage process increasesthe conversion of CO/H₂.

The reaction products comprise non-reacted CO and H₂, reaction productsmethanol, dimethyl ether, and CO₂, and slight amount of byproducts suchas alkane. Since the composition depends on the reaction rate andequilibrium of each of the reactions (1) through (3), the single stageprocess cannot increase the amount of solely a target product. Inparticular, methanol as an intermediate is unavoidably left in thereaction products.

The reaction rate in each reaction is controlled by changing the ratioof methanol synthesis catalyst, dehydration catalyst, and shift catalystThus the composition of the reaction products is controlled Since,however, these three types of reactions proceed simultaneously and sinceall of these three reactions are equilibrium reactions, the control hasa limitation owing to the limit of equilibrium. With that type ofreaction system and under a normal reaction condition, it is verydifficult to attain the selectivity of dimethyl ether over 95% (in theproducts excluding CO₂).

The difficulty is described below referring to a thermodynamiccalculation.

FIG. 15 shows reaction equilibria of H₂, CO, methanol, CO₂, and water,on the basis of reactions (1), (2), and (3). For instance, under acondition of 300° C. of reaction temperature, 50 atm of reactionpressure, and 1 of initial CO/H₂ ratio, the selectivity of dimethylether (starting material of CO, carbon molar basis, excluding CO₂) atthe reaction equilibrium is 98%. Establishing a reaction equilibrium is,however, impossible in actual process, and the selectivity of dimethylether becomes significantly lower level than the calculated value owingto the presence of methanol as an intermediate. In a state of a lowertemperature than the above example, for instance at 240° C. of reactiontemperature, 50 atm of reaction pressure, and 1 of initial CO/H₂ ratio,the selectivity of dimethyl ether at a reaction equilibrium becomes 99%which is somewhat higher than that in the above example. Under thecondition, however, the rate of methanol-synthesis reaction (1) is lowand actually the reaction equilibrium is never established

To cope with the situation, the method according to the presentinvention uses a mixture of a methanol synthesis catalyst, a dehydrationcatalyst or a dehydration and a shift catalyst in the first stagereaction to yield crude dimethyl ether, and uses a dehydration and/or ashift catalyst in the second stage reaction to convert most of theremaining methanol into dimethyl ether, thus attains a final highselectivity of dimethyl ether.

When the second stage reaction uses only a dehydration and/or a shiftcatalyst without applying a methanol synthesis catalyst, solely thedehydration and/or shift reaction proceeds and only these reactionsapproach the equilibrium state. In a reaction system where no methanolsynthesis catalyst exists, no additional methanol yields, and theremained methanol is converted into dimethyl ether, so the selectivityof dimethyl ether increases.

The above-described process, however, needs an additional reactor, whichincreases investment cost. In this respect, the inventors developed areactor having two separated sections therein using the similarity ofreaction condition (temperature and pressure) for both the first stagereaction and the second stage reaction. The reactor is described below.

The present invention was completed on the basis of the above-describedconcept, and the above-described object is attained by providing areaction apparatus for synthesizing dimethyl ether which comprises: areactor for synthesizing dimethyl ether from a mixed gas as a rawmaterial gas containing carbon monoxide and either or both of hydrogenand water vapor, or from a mixed gas as a raw material gas containingcarbon monoxide, either or both of hydrogen and water vapor, and carbondioxide, which reactor comprises a lower stage holding a catalyst slurryand an upper stage holding a catalyst fixed bed; a raw material gas feedpipe connected to the lower stage; a reaction product gas discharge pipeconnected to the upper stage.

The lower stage of the reactor according to the present invention is avertical type slurry-bed reactor which holds a catalyst slurry and has aheat exchanger therein to remove reaction heat.

The catalyst used in the slurry-bed is a combination of known methanolsynthesis catalyst, known dehydration catalyst, and known water gasshift catalyst. An applicable methanol synthesis catalyst includescommon industrial catalysts for methanol synthesis, such as those ofcopper oxide-zinc oxide system, zinc oxide-chromium oxide system, copperoxide-zinc oxide/chromium oxide system, copper oxide-zinc oxide/aluminasystem. Examples of dehydration catalyst are acid-base catalyst such asγ-alumina, silica, silica-alumina, and zeolite. Examples of the metallicoxide ingredient in zeolite are oxide of alkali metal such as sodium andpotassium, and oxide of alkali earth metal such as calcium andmagnesium. Examples of water gas shift catalyst are copper oxide-zincoxide system, copper oxide-chromium oxide-zinc oxide system, and ironoxide-chromium oxide system. Since methanol synthesis catalyst has astrong activity as a shift catalyst, it can substitute for the water gasshift catalyst A copper oxide supported by alumina is applicable as adehydration catalyst and also as a water gas shift catalyst.

The mixing ratio of the above-described methanol synthesis catalyst,dehydration catalyst, and water gas shift catalyst is not specificallylimited, and it depends on the kind of each ingredient and reactioncondition. Normally the ratio is often in an approximate range of from0.1 to 5 wt.parts of dehydration catalyst to 1 wt.parts of methanolsynthesis catalyst, more preferably in an approximate range of from 0.2to 2; in an approximate range of from 0.2 to 5 wt.parts of water gasshift catalyst, more preferably in an approximate range of from 0.5 to3. When the methanol synthesis catalyst substitutes for the water gasshift catalyst, the above-described amount of the water gas shiftcatalyst is added to the amount of the methanol synthesis catalyst.

The above-described catalyst is used in a powder state. A preferableaverage particle size is 300 μm or less, more preferably in anapproximate range of from 1 to 200 μm, and most preferably in anapproximate range of from 10 to 150 μm. To prepare powder of that rangeof particle size, the catalysts may further be pulverized.

An applicable kind of heating medium oil for dispersing catalyst isarbitrary if only the heating medium oil maintains liquid state underthe reaction condition. For example, the heating medium oil may behydrocarbons of aliphatic, aromatic, and alicyclic groups, alcohol,ether, ester, ketone, halide, or their mixture. Alternatively, gas oilafter removing sulfur ingredients, vacuum gas oil, high boiling pointdistillates of coal tar after being treated by hydrogenation,Fischer-Tropsch synthesis oil, and high boiling point food oil are alsoapplicable as the heating medium oil. The amount of catalyst in thesolvent depends on the kind of solvent and the reaction condition.Normally, a preferable range of the catalyst is from 1 to 50 wt. % tothe amount of solvent, more preferably in an approximate range of from10 to 30 wt. %.

A raw material gas is injected from a gas-injection nozzle mounted tothe slurry-bed reactor section to make the raw material gas contact withthe catalyst. With agitation to mix the slurry in the reactor, thereaction is enhanced. By inserting heat transfer tubes into the reactor,the reaction heat is removed from the reactor. The removed reaction heatmay be recovered to be utilized in other applications. Since aslurry-bed reactor is easy for agitation, the reaction heat is uniformlydispersed to the whole reaction system. As a result, hot spots hardlyappear, which makes the heat recovery easy. In addition, temperaturedistribution within the reactor hardly becomes irregular, so the amountof yielded byproducts is small. Furthermore, that type of reactor iseasy for charge and withdrawal of catalyst, and easy for start up andstop operation.

As described before, however, a reaction in a slurry-bed unavoidablyleave methanol in the reaction system as an intermediate product of thereaction.

To cope with the residual methanol issue, a fixed bed reactor section isinstalled at the upper part of the reactor to convert the remainingmethanol into dimethyl ether. A catalyst applied to the fixed bedreactor contains at least one of dehydration catalyst and dehydrationand shift catalyst As for the dehydration catalyst, one of theabove-described groups may be used. For the dehydration and shiftcatalyst, copper oxide supported by γ-alumina may be used as a catalysthaving both the dehydration activity and the shift activity. Thecatalyst in the fixed bed needs to be firmly held in the bed and alsoneeds to have a void that allows for a gas to pass freely through thefixed bed. To do this, the catalyst particle size may be granulated toan approximate size ranging from 1 to 20 mm, more preferably from 1.5 to10 mm. Members such as perforated plate and meshed material areinstalled at the bottom of the bed, as needed, to support the catalystparticles. The catalyst may be a porous block having lots of throughholepores.

The ratio of the amount of catalyst in a slurry-bed to that in a fixedbed depends on the activity of the catalyst in each bed. Normally, theweight ratio of the catalyst (slurry bed)/(fixed bed) is in a range offrom 1:10 to 10:1, more preferably from 1:5 to 5:1.

The boundary between the upper stage and the lower stage is in anarbitrary state if only gas freely passes therethrough. Entrainmentemitted from the lower fluidized catalyst layer is allowed to attach theupper fixed bed.

The raw material gas feed pipe is connected to the lower stage of thereactor to charge the raw material gas into the fluidized catalystlayer. The reaction product gas discharge pipe is connected to the upperstage to let the gas pass through the fixed catalyst bed and exit fromthe reactor.

Other sections than those described above may be similar with those inconventional reactors, and instruments such as a pressure gauge and athermometer, and if necessary, an agitator and an auxilliary rawmaterial charge line may be mounted to the reactor.

A dimethyl ether-synthesis apparatus that contains the dimethylether-synthesis reactor using the reactor according to the presentinvention may be the same configuration with a conventional apparatus.That is, at the exit of reactor, a series of equipment are connected ina sequent order of: a condenser to condense the heating medium oilvaporized from the reactor; a gas-liquid separator to separate thecondensed heating medium oil; a methanol-water separator which cools thereaction product gas to condense methanol and water to separate themfrom the gas phase; a non-reacted gas separator which further cools thegas to condense dimethyl ether and carbon dioxide, thus separatingthereof from carbon monoxide and hydrogen; and a CO₂ separator whichseparates dimethyl ether and carbon dioxide from the condensed mixture.Each of the methanol-water separator and the non-reacted gas separatormay further be divided into a condenser and a gas-liquid separator.Alternatively, methanol, water, dimethyl ether, and carbon dioxide maybe condensed or solidified together, then carbon monoxide and hydrogenmay be separated, followed by separating dimethyl ether from thecondensed solid.

FIG. 16 shows an example of the apparatus implementing theabove-described method. The apparatus comprises a reactor R, amethanol-water separator S1, a non-reacted gas separator S2, and a CO₂separator S3. A raw material gas line 502 is connected to the bottom ofreactor R. To the raw material gas line 502, a make up (fresh) gas line501 which supplies fresh raw material gas and a recycle gas line 507through which the non-reacted CO and H2 are recycled to charge areconnected. A reaction product gas line 503 to discharge the reactionproducts connects the top of reactor R to the inlet of methanol-waterseparator S1, (a condenser for heating medium oil and a gas-liquidseparator therefor are installed upstream of the methanol-waterseparator S1, at need). A methanol-water line 504 is connected to theexit of methanol-water on the methanol-water separator S1. A reactionproduct gas line 505 is connected to the exit of reaction products onthe methanol-water separator S1. The other end of the reaction productgas line 505 is connected to the inlet of the non-reacted gas separatorS2. The other end of the recycle gas line 507 is connected to the exitof non-reacted gas on the non-reacted gas separator S2. The recycle gasline 507 is provided with a branched purge line 510 to withdraw apart ofthe gas. A DME, CO₂ line 506 is connected to the exit of DME, CO₂ on thenon-reacted gas separator S2. Other end of the DME, CO₂ line 6 isconnected to the CO₂ separator S3. A CO₂ line 508 is connected to theCO₂ exit on the CO₂, separator S3. A DME line 509 is connected to theexit of DME on the CO₂ separator S3.

The mixing ratio of hydrogen and carbon monoxide may be in a range offrom 20 to 0.1 as H₂/CO molar ratio, more preferably in a range of from10 to 0.2. In the case of a mixed gas with a significantly low ratio of(H₂/CO), for example, 0.1 or less, or in the case of solely carbonmonoxide without containing hydrogen, it is necessary to separatelysupply steam to conduct the shift reaction in the reactor to convert apart of the carbon monoxide into hydrogen and carbon dioxide. Apreferable charge rate of steam is 1 or less to the charge rate of CO. Apreferable amount of carbon dioxide yielded from the reaction is 50% orless.

A preferable condition for both the upper and lower stage reactions isthe reaction temperature in a range of from 150 to 400° C., particularlyin a range of from 200 to 350° C. The reaction temperature below 150° C.and above 400° C. results in a reduction of carbon monoxide conversion.A preferable reaction pressure is in a range of from 10 to 300 kg/cm²,particularly in a range of from 15 to 70 kg/cm². The reaction pressurebelow 10 kg/cm² results in a low conversion of carbon monoxide, and thatabove 300 kg/cm² requires a special design of reactor and isuneconomical because of the need of a large amount of energy forpressurizing the system. A preferable space velocity (charge rate ofmixed gas per 1 g of catalyst under standard condition) is in a range offrom 100 to 50000 l/kg·h, and particularly preferable from 500 to 30000l/kg·h. The space velocity above 50000 l/kg·h degrades the conversion ofcarbon monoxide, and that below 100 l/kg·h is uneconomical because ofthe need of an excessively large reactor.

For the reactor according to the present invention, the use of acombination of methanol synthesis catalyst, dehydration catalyst, andshift catalyst in the first stage reaction increases the conversion ofCO/H₂ in the raw material gas, and the use of dehydration catalystand/or shift catalyst in the second stage reaction converts most part ofthe remained methanol into dimethyl ether, thus the selectivity ofdimethyl ether increases. In spite of two stage reaction, the reactor isdesigned to a compact type configuring two separated sections integratedin a single vessel utilizing almost the same reaction condition(temperature and pressure) for both stages.

-   Amount of CO gas charged to reactor (Nl/min.):Fin(CO)-   Amount of CO gas discharged from reactor:Fout(CO)-   Amount of DME gas discharged from reactor:Fout(DME)-   Amount of MeOH gas discharged from reactor:Fout(MeHO)-   Amount of methane gas discharged from reactor:Fout(CH₄)

$\begin{matrix}{( {{CO}\mspace{14mu}{conversion}} ) = \frac{{{Fin}({CO})} - {{Fout}({CO})}}{{Fin}({CO})}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{14mu}{dimethylether}} ) = \frac{2 \times {{Fout}({DME})}}{\begin{matrix}{{2 \times {{Fout}({DME})}} +} \\{{{Fout}({MeOH})} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}} \\{( {{Selectivity}\mspace{14mu}{of}\mspace{14mu}{methanol}} ) = \frac{{Fout}({MeOH})}{\begin{matrix}{2 \times {{{Fout}({DME})} \div}} \\{{{Fout}( {{MeO}H} )} +} \\{{Fout}( {CH}_{4} )}\end{matrix}}}\end{matrix} \times 100(\%)$

EXAMPLE

Catalyst A: CuO—ZnO—Al₂O₃ catalyst.

Each of 185 g of copper nitrate (Cu(NO₃)₂3H₂O), 117 g of zinc nitrate(Zn(NO₃)₂6H₂O), and 52 g of aluminum nitrate (Al(NO₃)₃9H₂O) weredissolved into about 1 liter of ion-exchanged water. Separately, about1.4 kg of sodium carbonate (Na₂CO₃) was dissolved into about 1 liter ofion-exchanged water. Both of the solutions were added dropwise to about3 liters of ion-exchanged water in a stainless steel vessel which wascontrolled at about 60° C. within about 2 hours. while maintaining thecontents to pH 7.0±0.5. Then, the contents were allowed to stand forabout 1 hour for aging. When, during the treatment, the pH value wentout from a range of pH 7.0±0.5, an aqueous solution of about 1mole/liter sodium carbonate was added dropwise to keep the range of pH7.0±0.5. The resultant precipitate was filtered, and the cake was rinsedby ion-exchanged water until nitric acid ion was not detected anymore.After the rinse, the cake was dried at 120° C. for 24 hours, followed bycalcining thereof in air at 350° C. for 5 hours to obtain the targetcatalyst. Analysis of the thus obtained catalyst gave the composition asCuO:ZnO:Al₂O₃=61:32:7 (by weight).

Catalyst B: CuO—Al₂O₃ catalyst

A 15.7 g of copper acetate (Cu(CH₃COO)₂H₂O) was dissolved into about 200ml of ion-exchanged water. A 95 g of γ-alumina (N612, Nikki Kagaku Co.)was further added to the mixture. The mixture was then vaporized to dry.The dried material was calcined in air at 450° C. for 4 hours. Thecalcined material was treated in hydrogen gas stream at 400° C. for 3hours to obtain a catalyst Analysis of the catalyst gave the compositionas Cu:Al₂O₃=5:95 (by weight).

Each of the catalyst thus prepared was pulverized in a ball mill to aparticle size of 120 μm or less.

The lower stage of the reactor was filled with 5584 ml of n-hexadecaneas the heating medium oil, 430 g of the catalyst A, and 215 g ofcatalyst B: that is, (catalyst A/catalyst B)=2/1, and (catalyst/heatingmedium oil)=15/100. The upper stage of the reactor was filled solelywith 645 g of catalyst B.

(Preliminary Reduction)

Under a condition of 10 kg/cm² of reactor pressure, 220° C. of reactortemperature, a mixed gas (H₂/N₂=1/4) was introduced to the reactor at aflow rate of 10 l/min. for 12 hours to conduct preliminary reduction.

A gas of H₂/CO=1/1 was introduced to the reaction system at a flow rateof 18 l/min. to conduct the dimethyl ether synthesis under a conditionof 50 kg/cm²G and 260° C. for both the upper and the lower stages. Gasanalysis was conducted by gas-chromatography, and the gas flow rate atthe exit of reaction system was determined by a gas meter. From theanalysis and flow rate determination, CO conversion and selectivity ofeach reaction product were calculated, (carbon molar basis, excludingCO₂). The result showed 41.0% of CO conversion, 95.5% of DMEselectivity, 4.4% of methanol selectivity, and 0.1% of methaneselectivity.

Comparative Example

Dimethyl ether synthesis was conducted under the same conditions in thepreceding Example except that the upper stage of the reactor did notcontain catalyst The result showed 34.0% of CO conversion, 67.1% of DMEselectivity, 32.8% of methanol selectivity, and 0.1% of methaneselectivity. The DME selectivity was at a low level.

The reactor according to the present invention provides dimethyl etherfrom carbon monoxide and hydrogen (or water vapor) at a high conversionand high selectivity. As a result, high purity dimethyl ether isobtained from the reaction products, thus allowing mass production ofdimethyl ether at a low cost.

1. A catalyst for producing dimethyl ether, the catalyst being producedby a method comprising forming a layer comprising a methanol synthesiscatalyst which comprises copper oxide, zinc oxide and alumina, aroundalumina particles, the alumina particles having an average size of 200μm or less, wherein the methanol synthesis catalyst layer formed aroundsaid alumina particles is in an amount of 0.05 to 5 parts by weight to 1part by weight of the alumina, wherein the layer is produced by a methodcomprising: (a) forming a slurry by introducing the alumina particlesinto an aqueous solution containing a metallic salt of one or moreactive elements of the methanol synthesis catalyst; (b) heating theslurry at a temperature of 50 to 90° C. to provide a heated slurry; (c)neutralizing the heated slurry with a base solution, whereby the one ormore active elements of the methanol synthesis catalyst form deposits onthe alumina particles; (d) allowing the slurry to stand for a sufficienttime or subjection the slurry to a mild agitation for aging tosufficiently develop the deposits; and (e) washing the layer with anacidic aqueous solution.
 2. The catalyst of claim 1, wherein the averagesize of the alumina particles is 1 to 100 μm.
 3. The catalyst of claim2, wherein the average size of the alumina particles is 1 to 50 μm. 4.The catalyst of claim 3, wherein the methanol synthesis catalyst has aweight ratio of the copper oxide: the zinc oxide: the alumina being 1:(0.05 to 20) : (up to 2).
 5. The catalyst of claim 3, wherein the acidicaqueous solution is an aqueous solution of an inorganic acid.
 6. Thecatalyst of claim 3, wherein the acidic aqueous solution is an aqueoussolution of an organic acid.
 7. The catalyst of claim 2, wherein themethanol synthesis catalyst has a weight ratio of the copper oxide: thezinc oxide: the alumina being 1: (0.05 to 20) : (up to 2).
 8. Thecatalyst of claim 2, wherein the acidic aqueous solution is an aqueoussolution of an inorganic acid.
 9. The catalyst of claim 2, wherein theacidic aqueous solution is an aqueous solution of an organic acid. 10.The catalyst of claim 1, wherein the methanol synthesis catalyst has aweight ratio of the copper oxide: the zinc oxide: the alumina being 1:(0.05 to 20) : (up to 2).
 11. The catalyst of claim 10, wherein theacidic aqueous solution is an aqueous solution of an inorganic acid. 12.The catalyst of claim 10, wherein the acidic aqueous solution is anaqueous solution of an organic acid.
 13. The catalyst of claim 1,wherein the acidic aqueous solution is an aqueous solution of aninorganic acid.
 14. The catalyst of claim 13, wherein the inorganic acidis selected from the group consisting of nitric acid and hydrochloricacid.
 15. The catalyst of claim 1, wherein the acidic aqueous solutionis an aqueous solution of an organic acid.
 16. The catalyst of claim 15,wherein the organic acid is acetic acid.
 17. The catalyst of claim 2,wherein the slurry in step (b) is heated to a temperature of 60 to 85°C.
 18. The catalyst of claim 1, wherein the methanol synthesis layerformed around the alumina particle is in an amount of 0.1 to 3 parts byweight to 1 part by weight of the alumina.
 19. The catalyst of claim 1,wherein the methanol synthesis layer formed around the alumina particleis in an amount of 0.5 to 2 parts by weight to 1 part by weight of thealumina.
 20. A catalyst for producing dimethyl ether, the catalyst beingproduced by a method comprising forming a layer comprising a methanolsynthesis catalyst which comprises chromium oxide, zinc oxide andalumina, around alumina particles, the alumina particles having anaverage size of 200 μm or less, wherein the methanol synthesis catalystlayer formed around said alumina particles is in an amount of 0.05 to 5parts by weight to 1 part by weight of the alumina, wherein the layer isproduced by a method comprising: (a) forming a slurry by introducing thealumina particles into an aqueous solution containing a metallic salt ofone or more active elements of the methanol synthesis catalyst; (b)heating the slurry at a temperature of 50 to 90° C. to provide a heatedslurry; (c) neutralizing the heated slurry with a base solution, wherebythe one or more active elements of the methanol synthesis catalyst formdeposits on the alumina particles; (d) allowing the slurry to stand fora sufficient time or subjecting the slurry to a mild agitation for agingto sufficiently develop the deposits; and (e) washing the layer with anacidic aqueous solution.
 21. The catalysts of claim 20, wherein theaverage size of the alumina particles is 1 to 100 μm.
 22. The catalystof claim 21, wherein the methanol synthesis catalyst has a weight ratioof the zinc oxide: the chromium oxide: the alumina being 1: (0.1 to 10):(up to 2).
 23. The catalyst of claim 21, wherein the acidic aqueoussolution is an aqueous solution of an inorganic acid.
 24. The catalystof claim 21, wherein the acidic aqueous solution is an aqueous solutionof an organic acid.
 25. The catalyst of claim 20, wherein the averagesize of the alumina particles is 1 to 50 μm.
 26. The catalyst of claim25, wherein the methanol synthesis catalyst has a weight ratio of thezinc oxide: the chromium oxide: the alumina being 1: (0.1 to 10): (up to2).
 27. The catalyst of claim 25, wherein the acidic aqueous solution isan aqueous solution of an inorganic acid.
 28. The catalyst of claim 25,wherein the acidic aqueous solution is an aqueous solution of an organicacid.
 29. The catalyst of claim 20, wherein the methanol synthesiscatalyst has a weight ratio of the zinc oxide: the chromium oxide: thealumina being 1: (0.1 to 10): (up to 2).
 30. The catalyst of claim 29,wherein the acidic aqueous solution is an aqueous solution of aninorganic acid.
 31. The catalyst of claim 29, wherein the acidic aqueoussolution is an aqueous solution of an organic acid.
 32. The catalyst ofclaim 20, wherein the acidic aqueous solution is an aqueous solution ofan inorganic acid.
 33. The catalyst of claim 32, wherein the inorganicacid is selected from the group consisting of nitric acid andhydrochloric acid.
 34. The catalyst of claim 20, wherein the acidicaqueous solution in an aqueous solution of an inorganic acid.
 35. Thecatalyst of claim 34, wherein the organic acid is acetic acid.
 36. Thecatalyst of claim 20, wherein the slurry in step (b) is heated to atemperature of 60 to 85° C.
 37. A catalyst for producing dimethyl ether,the catalyst being produced by a method comprising forming a layercomprising a methanol synthesis catalyst which comprises chromium oxide,copper oxide, zinc oxide and alumina, around alumina particles, thealumina particles having an average size of 200μ or less, wherein themethanol synthesis catalyst layer formed around said alumina particlesis in an amount of 0.05 to parts by weight to 1 part by weight of thealumina, wherein the layer is produced by a method comprising: (a)forming a slurry by introducing the alumina particles into an aqueoussolution containing a metallic salt of one of more active elements ofthe methanol synthesis catalyst; (b) heating the slurry at a temperatureof 50 to 90° C. to provide a heated slurry; (c) neutralizing the heatedslurry with a base solution, whereby the one or more active elements ofthe methanol synthesis catalyst form deposits on the alumina particles;(d) allowing the slurry to stand for a sufficient time or subjecting theslurry to a mild agitation for aging to sufficiently develop thedeposits; and (e) washing the layer with an acidic aqueous solution. 38.The catalyst of claim 37, wherein the acidic aqueous solution is anaqueous solution of an inorganic acid.
 39. The catalyst of claim 38,wherein the inorganic acid is selected from the group consisting ofnitric acid and hydrochloric acid.
 40. The catalyst of claim 37, whereinthe acidic aqueous solution is an aqueous solution of an organic acid.41. The catalyst of claim 40, wherein the organic acid is acetic acid.42. The catalyst of claim 37, wherein the slurry in step (b) is heatedto a temperature of 60 to 85° C.
 43. A method for producing a catalystcomprising forming a layer comprising a methanol synthesis catalystwhich comprises copper oxide, zinc oxide and alumina, around aluminaparticles, the alumina particles having an average size of 200 μm orless, wherein the methanol synthesis catalyst layer formed around saidalumina particles is in an amount of 0.05 to 5 parts by weight to 1 partby weight of the alumina, wherein the layer is produced by a methodcomprising: (a) forming a slurry by introducing the alumina particlesinto an aqueous solution containing a metallic salt of one or moreactive elements of the methanol synthesis catalyst; (b) heating theslurry at a temperature of 50 to 90° C. to provide a heated slurry; (c)neutralizing the heated slurry with a base solution, whereby the one ormore active elements of the methanol synthesis catalyst from deposits onthe alumina particles; (d) allowing the slurry to stand for a sufficienttime of subjecting the slurry to mild agitation for aging tosufficiently develop the deposits; and (e) washing the layer with anacidic aqueous solution.